Reaction systems for making N-(phosphonomethyl) glycine compounds

ABSTRACT

This invention generally relates to liquid phase oxidation processes for making N-(phosphonomethyl)glycine (also known in the agricultural chemical industry as glyphosate) and related compounds. This invention, for example, particularly relates to processes wherein an N-(phosphonomethyl)iminodiacetic acid (NPMIDA) substrate (i.e., N-(phosphonomethyl)iminodiacetic acid, a salt of N-(phosphonomethyl)iminodiacetic acid, or an ester of N-(phosphonomethyl)iminodiacetic acid) is continuously oxidized to form an N-(phosphonomethyl)glycine product (i.e., N-(phosphonomethyl)glycine, a salt of N-(phosphonomethyl)glycine, or an ester of N-(phosphonomethyl)glycine). This invention also, for example, particularly relates to processes wherein an N-(phosphonomethyl)iminodiacetic acid substrate is oxidized to form an N-(phosphonomethyl)glycine product, which, in turn, is crystallized (at least in part) in an adiabatic crystallizer.

This application is a divisional of U.S. patent application Ser. No.09/863,885, filed May 22, 2001, now U.S. Pat. No. 7,015,351, whichclaims the benefit of U.S. provisional application Ser. No. 60/206,562,filed May 22, 2000, U.S. provisional application Ser. No. 60/220,140,filed Jul. 21, 2000, and U.S. provisional application Ser. No.60/230,240, filed Sep. 1, 2000, the entire disclosures of which areincorporated herein by reference.

FIELD OF THE INVENTION

This invention generally relates to liquid phase oxidation processes formaking N-(phosphonomethyl)glycine (also known in the agriculturalchemical industry as glyphosate) and related compounds. This invention,for example, particularly relates to processes wherein anN-(phosphonomethyl)iminodiacetic acid (NPMIDA) substrate (i.e.,N-(phosphonomethyl)iminodiacetic acid, a salt ofN-(phosphonomethyl)iminodiacetic acid, or an ester ofN-(phosphonomethyl)iminodiacetic acid) is continuously oxidized to forman N-(phosphonomethyl)glycine product (i.e., N-(phosphonomethyl)glycine,a salt of N-(phosphonomethyl)glycine, or an ester ofN-(phosphonomethyl)glycine). This invention also, for example,particularly relates to processes wherein anN-(phosphonomethyl)iminodiacetic acid substrate is oxidized to form anN-(phosphonomethyl)glycine product, which, in turn, is crystallized (atleast in part) in an adiabatic crystallizer.

BACKGROUND OF THE INVENTION

N-(phosphonomethyl)glycine is described by Franz in U.S. Pat. No.3,799,758. N-(phosphonomethyl)glycine and its salts are convenientlyapplied as a post-emergent herbicide in an aqueous formulation. It is ahighly effective and commercially important broad-spectrum herbicideuseful in killing or controlling the growth of a wide variety of plants,including germinating seeds, emerging seedlings, maturing andestablished woody and herbaceous vegetation, and aquatic plants.

One of the more widely accepted methods of makingN-(phosphonomethyl)glycine compounds comprises oxidatively cleaving acarboxymethyl substituent from an N-(phosphonomethyl)iminodiacetic acidsubstrate. Over the years, a wide variety of methods have been disclosedfor conducting this oxidation. See generally, Franz, et al., Glyphosate:A Unique Global Herbicide (ACS Monograph 189, 1997) at pp. 233-62 (andreferences cited therein); Franz (U.S. Pat. No. 3,950,402); Hershman(U.S. Pat. No. 3,969,398); Chou (U.S. Pat. No. 4,624,937); Chou (U.S.Pat. No. 4,696,772); Ramon et al. (U.S. Pat. No. 5,179,228); Felthouse(U.S. Pat. No. 4,582,650); Siebenhaar et al. (PCT/EP99/04587); and Ebneret al. (International Publication No. WO 99/43430). Although many ofthese processes produce suitable yields of variousN-(phosphonomethyl)glycine products, a need continues to exist for animproved process for oxidizing N-(phosphonomethyl)iminodiacetic acidsubstrates. Desirable improvements include increased throughput, reducedcost per unit of N-(phosphonomethyl)glycine product, and reducedconcentrations of undesirable by-products (e.g., formaldehyde, formicacid, N-methyl-N-(phosphonomethyl)glycine (NMG), andaminomethylphosphonic acid (AMPA)).

SUMMARY OF THE INVENTION

This invention provides, in part, for economical processes for oxidizingN-(phosphonomethyl)iminodiacetic acid, salts ofN-(phosphonomethyl)iminodiacetic acid, and esters ofN-(phosphonomethyl)iminodiacetic acid to formN-(phosphonomethyl)glycine, salts of N-(phosphonomethyl)glycine, andesters of N-(phosphonomethyl)glycine. This invention also provideseffective methods for purifying and/or concentrating theN-(phosphonomethyl)glycine product obtained in the oxidation reactionmixture.

Briefly, therefore, the present invention is directed to a process formaking an N-(phosphonomethyl)glycine product. The process comprisesintroducing an aqueous feed stream comprising anN-(phosphonomethyl)iminodiacetic acid substrate into an oxidationreactor system in which the N-(phosphonomethyl)iminodiacetic acidsubstrate is oxidized in the presence of an oxidation catalyst toproduce a reaction product solution comprisingN-(phosphonomethyl)glycine product. The reaction product solution isdivided into plural fractions comprising a primary fraction and asecondary fraction. N-(phosphonomethyl)glycine product crystals areprecipitated from the primary fraction to produce a primary productslurry comprising precipitated N-(phosphonomethyl)glycine productcrystals and a primary mother liquor, while N-(phosphonomethyl)glycineproduct crystals are also precipitated from an aqueous secondarycrystallization feed mixture comprising N-(phosphonomethyl)glycineproduct contained in the secondary fraction to produce a secondaryproduct slurry comprising precipitated N-(phosphonomethyl)glycineproduct crystals and a secondary mother liquor.

In another embodiment, the process for making anN-(phosphonomethyl)glycine product comprises introducing an aqueous feedstream comprising an N-(phosphonomethyl)iminodiacetic acid substrateinto an oxidation reactor system and oxidizing theN-(phosphonomethyl)iminodiacetic acid substrate in the oxidation reactorsystem in the presence of an oxidation catalyst to produce a reactionproduct solution containing N-(phosphonomethyl)glycine product.N-(phosphonomethyl)glycine product crystals are precipitated from thereaction product solution to produce a primary product slurry comprisingprecipitated N-(phosphonomethyl)glycine product crystals and a primarymother liquor. Water is then evaporated from the primary mother liquorto thereby precipitate additional N-(phosphonomethyl)glycine productcrystals and produce a secondary mother liquor.

In another embodiment, the process for making anN-(phosphonomethyl)glycine product comprises introducing an aqueous feedstream comprising an N-(phosphonomethyl)iminodiacetic acid substrateinto a primary oxidation reactor system comprising one or more oxidationreaction zones. The N-(phosphonomethyl)iminodiacetic acid substrate isoxidized in the primary oxidation reactor system to produce a reactionproduct solution comprising N-(phosphonomethyl)glycine product andunreacted N-(phosphonomethyl)iminodiacetic acid substrate. The reactionproduct solution is divided into plural fractions comprising a primaryfraction and a secondary oxidation reactor feed fraction.N-(phosphonomethyl)glycine product crystals are precipitated from theprimary fraction to produce a primary product slurry comprisingprecipitated N-(phosphonomethyl)glycine product crystals and a primarymother liquor. The secondary oxidation reactor feed fraction isintroduced into a secondary oxidation reactor system comprising one ormore oxidation reaction zones. The N-(phosphonomethyl)iminodiacetic acidsubstrate is oxidized in the secondary oxidation reactor system toproduce a secondary oxidation reactor effluent comprisingN-(phosphonomethyl)glycine product. Thereafter,N-(phosphonomethyl)glycine product crystals are precipitated from thesecondary oxidation reactor effluent to produce a secondary productslurry comprising precipitated N-(phosphonomethyl)glycine productcrystals and a secondary mother liquor.

The present invention is also directed to a process for preparing anN-(phosphonomethyl)glycine product by oxidizing anN-(phosphonomethyl)iminodiacetic acid substrate. The process comprisesintroducing the N-(phosphonomethyl)iminodiacetic acid substrate into aliquid reaction medium comprising the N-(phosphonomethyl)glycine productwithin an oxidation reaction zone. The oxidation reaction zone issubstantially back-mixed in the liquid phase and contains a catalyst forthe oxidation reaction in contact with the liquid reaction medium. Anoxidizing agent is also introduced into the oxidation reaction zonewherein the N-(phosphonomethyl)iminodiacetic acid substrate iscontinuously oxidized to form the N-(phosphonomethyl)glycine product. Areaction mixture effluent comprising the N-(phosphonomethyl)glycineproduct is continuously withdrawn from the oxidation reaction zone.

In another embodiment, the process comprises introducing theN-(phosphonomethyl)iminodiacetic acid substrate into a liquid reactionmedium within an oxidation reaction zone. The liquid reaction mediumcomprises the N-(phosphonomethyl)glycine product and has a particulateheterogeneous catalyst for the oxidation reaction suspended therein. Anoxidizing agent is also introduced into the oxidation reaction zonewherein the N-(phosphonomethyl)iminodiacetic acid substrate iscontinuously oxidized in the liquid reaction medium to form theN-(phosphonomethyl)glycine product. A reaction mixture effluentcomprising the N-(phosphonomethyl)glycine product is continuouslywithdrawn from said oxidation reaction zone. The particulate catalyst iscontinuously separated from the reaction mixture effluent to form acatalyst recycle stream comprising the separated catalyst. At least aportion of the particulate catalyst contained in the catalyst recyclestream is introduced into said oxidation reaction zone.

The present invention is further directed to a continuous process forpreparing an N-(phosphonomethyl)glycine product by oxidizing anN-(phosphonomethyl)iminodiacetic acid substrate in a reactor system. Theprocess comprises introducing an aqueous feed stream comprising theN-(phosphonomethyl)iminodiacetic acid substrate and an oxidizing agentinto a first oxidation reaction zone. TheN-(phosphonomethyl)iminodiacetic acid substrate is continuously oxidizedin the first oxidation reaction zone to form theN-(phosphonomethyl)glycine product. An intermediate reaction mixtureeffluent comprising the N-(phosphonomethyl)glycine product and unreactedN-(phosphonomethyl)iminodiacetic acid substrate is continuouslywithdrawn from the first oxidation reaction zone. An intermediateaqueous feed stream comprising N-(phosphonomethyl)glycine product andunreacted N-(phosphonomethyl)iminodiacetic acid substrate obtained inthe intermediate reaction mixture effluent is continuously introducedinto a second oxidation reaction zone along with an oxidizing agentwherein N-(phosphonomethyl)iminodiacetic acid substrate is continuouslyoxidized to form additional N-(phosphonomethyl)glycine product. Areaction mixture effluent comprising the N-(phosphonomethyl)glycineproduct is continuously withdrawn from the second oxidation reactionzone.

The present invention is also directed to processes for concentratingand recovering the N-(phosphonomethyl)glycine product. In oneembodiment, a process for removing water from an aqueous startingsolution comprising N-(phosphonomethyl)glycine product and crystallizingN-(phosphonomethyl)glycine product therefrom is provided. The processcomprises introducing an aqueous evaporation feed mixture comprising theaqueous starting solution into an evaporation zone. Water is evaporatedfrom the feed mixture in the evaporation zone in the presence of solidparticulate N-(phosphonomethyl)glycine product, thereby producing avapor phase comprising water vapor, precipitatingN-(phosphonomethyl)glycine product from the aqueous liquid phase, andproducing an evaporation product comprising N-(phosphonomethyl)glycineproduct solids and a mother liquor that is substantially saturated orsupersaturated in N-(phosphonomethyl)glycine product. A ratio ofparticulate N-(phosphonomethyl)glycine product solids to mother liquoris maintained in the evaporation zone which exceeds the ratio ofN-(phosphonomethyl)glycine product solids incrementally produced by theeffects of evaporation to mother liquor incrementally produced thereby.

In a further embodiment, the process comprises introducing anevaporation feed mixture comprising the aqueous starting solution into avapor/liquid separation zone wherein the pressure is below the vaporpressure of the mixture. This allows water to flash from the evaporationfeed mixture, producing a vapor phase comprising water vapor andincreasing the concentration of N-(phosphonomethyl)glycine product inthe remaining liquid phase to a concentration in excess of thesolubility of N-(phosphonomethyl)glycine product. As a result,N-(phosphonomethyl)glycine product precipitates from the liquid phase toproduce a first slurry stream comprising particulateN-(phosphonomethyl)glycine product in a saturated or supersaturatedmother liquor. The vapor phase is separated from the first slurry streamand the first slurry stream is introduced into a decantation zone inwhich a supernatant liquid comprising a fraction of the mother liquor isseparated from a second slurry stream comprising precipitatedN-(phosphonomethyl)glycine product and mother liquor. The decantationzone has an inlet for the first slurry, a decantation liquid exit forthe supernatant liquid spaced above the inlet, and an exit for thesecond slurry vertically spaced above the inlet but below thesupernatant liquid exit. The relative rates at which the first slurry isintroduced into the decantation zone, the second slurry is drawn offthrough the second slurry exit and the supernatant liquid is drawn offthrough the decantation liquid exit are maintained such that the upwardflow velocity in a lower region of the decantation zone below the secondslurry exit is sufficient to maintain precipitatedN-(phosphonomethyl)glycine product in suspension (i.e., entrained) inthe liquid phase while the upward flow velocity in an upper region ofthe decantation zone above the second slurry exit is below thesedimentation velocity of at least 80% by weight of theN-(phosphonomethyl)glycine product particles in the lower region.

In a still further embodiment, the process comprises introducing anaqueous evaporation feed mixture comprising the aqueous startingsolution into an evaporation zone. Water is evaporated from the feedmixture in the evaporation zone in the presence of solid particulateN-(phosphonomethyl)glycine product, thereby producing a vapor phasecomprising water vapor, precipitating N-(phosphonomethyl)glycine productfrom the aqueous liquid phase, and producing an evaporation productcomprising N-(phosphonomethyl)glycine product solids and a mother liquorthat is substantially saturated or supersaturated inN-(phosphonomethyl)glycine product. The evaporation product is dividedto provide an N-(phosphonomethyl)glycine product solids fraction that isrelatively depleted in mother liquor and a mother liquor fraction thatis relatively depleted in N-(phosphonomethyl)glycine product solids. Aratio of particulate N-(phosphonomethyl)glycine product solids to motherliquor is maintained in the evaporation zone that exceeds the ratio ofN-(phosphonomethyl)glycine product solids incrementally produced by theeffects of evaporation to mother liquor incrementally produced thereby.

The present invention is also directed to integrated processes for thepreparation of an oxidation reaction mixture effluent comprising theN-(phosphonomethyl)glycine product and thereafter concentrating andrecovering the product. In one embodiment, the process comprisesintroducing an aqueous feed mixture comprisingN-(phosphonomethyl)iminodiacetic acid substrate into a liquid reactionmedium and catalytically oxidizing the N-(phosphonomethyl)iminodiaceticacid substrate in the aqueous liquid reaction medium thereby producingan oxidation reaction mixture comprising N-(phosphonomethyl)glycineproduct. A primary crystallization feed mixture comprisingN-(phosphonomethyl)glycine product produced in the reaction mixture iscooled, thereby precipitating N-(phosphonomethyl)glycine product andproducing a primary mother liquor comprising N-(phosphonomethyl)glycineproduct. After separating precipitated N-(phosphonomethyl)glycineproduct from the primary mother liquor, primary mother liquor isrecycled and introduced into the liquid reaction medium whereinN-(phosphonomethyl)iminodiacetic acid substrate is oxidized toN-(phosphonomethyl)glycine product.

In a further embodiment of the present invention, the process comprisesintroducing an aqueous feed mixture comprising anN-(phosphonomethyl)iminodiacetic acid substrate into a catalytic reactorsystem comprising one or more catalytic reaction zones. TheN-(phosphonomethyl)iminodiacetic acid substrate is catalyticallyoxidized to N-(phosphonomethyl)glycine product in the catalytic reactorsystem to produce a product mixture which is then divided into a primaryfraction and a secondary fraction. N-(phosphonomethyl)glycine productfrom the primary fraction is crystallized to produce a solidN-(phosphonomethyl)glycine product fraction and a primary mother liquor.Primary mother liquor is recycled for use as a source of water in thepreparation of the feed mixture introduced into the catalytic reactorsystem.

The present invention is further directed to a continuous processes forthe catalytic oxidation of an N-(phosphonomethyl)iminodiacetic acidsubstrate to produce an N-(phosphonomethyl)glycine product. In oneembodiment, the process comprises introducing a liquid phase feed streamcomprising an aqueous feed stream comprising theN-(phosphonomethyl)iminodiacetic acid substrate into a primary oxidationreaction zone, the primary oxidation reaction zone comprising a primaryfixed bed containing an oxidation catalyst. An oxidizing agent isintroduced into the primary oxidation reaction zone wherein theN-(phosphonomethyl)iminodiacetic acid substrate is continuously oxidizedto the N-(phosphonomethyl)glycine product, thereby producing a primaryreaction mixture comprising the N-(phosphonomethyl)glycine product andunreacted N-(phosphonomethyl)iminodiacetic acid substrate. The primaryreaction mixture is withdrawn from the primary oxidation reaction zone.The difference in unit weight sensible heat content between the reactionmixture and the aqueous feed stream is maintained less than theexothermic reaction heat generated in the reaction zone per unit weightof the aqueous feed stream.

In another embodiment, the continuous process comprises introducing anaqueous feed stream comprising the N-(phosphonomethyl)iminodiacetic acidsubstrate into the first of a series of oxidation reaction zones, eachof the series of oxidation reaction zones comprising an oxidationcatalyst. The N-(phosphonomethyl)iminodiacetic acid substrate isoxidized in the first oxidation reaction zones to produce anintermediate oxidation reaction product. The intermediate oxidationreaction product is introduced into a second oxidation reaction zonecomprising a fixed bed containing a noble metal on carbon catalyst,wherein by-product formaldehyde and/or formic acid is oxidized.

In a further embodiment, the continuous process comprises introducing afirst component feed stream comprising anN-(phosphonomethyl)iminodiacetic acid substrate into the first of aseries of continuous reaction zones, each of the series of reactionzones comprising an oxidation catalyst. An oxidant is introduced intothe first of the series of reaction zones wherein the substrate iscatalytically oxidized to produce an intermediate reaction mixturestream containing N-(phosphonomethyl)glycine product. The intermediatereaction mixture exiting the first reaction zone is transferred to thesecond of the series of reaction zones wherein the substrate iscatalytically oxidized. An intermediate reaction mixture is withdrawnfrom each of the reaction zones and introduced into each succeedingreaction zone. An additional component feed stream is introduced intoeach of one or more of the reaction zones succeeding the first reactionzone in the series, each the additional feed stream comprising anN-(phosphonomethyl)iminodiacetic acid substrate. An oxidant isintroduced into one or more of the reaction zones succeeding the firstreaction zone in the series. A final reaction product is withdrawn fromthe last in the series of reaction zones.

In a further embodiment, the continuous process comprises introducing anaqueous feed stream comprising the N-(phosphonomethyl)iminodiacetic acidsubstrate into an oxidation reaction zone comprising a fixed bedcontaining an oxidation catalyst. An O₂-containing gas is introducedinto the oxidation reaction zone wherein theN-(phosphonomethyl)iminodiacetic acid substrate is continuously oxidizedto the N-(phosphonomethyl)glycine product, thereby producing anoxidation reaction mixture comprising the N-(phosphonomethyl)glycineproduct. The ratio of the mass flow rate of the liquid phase to the massflow rate of gas phase in the fixed bed is between about 20 and about800.

In a further embodiment, the continuous process comprises introducing anaqueous feed stream comprising the N-(phosphonomethyl)iminodiacetic acidsubstrate into an oxidation reaction zone comprising a fixed bedcontaining an oxidation catalyst. An O₂-containing gas is introducedinto the oxidation reaction zone wherein theN-(phosphonomethyl)iminodiacetic acid substrate is continuously oxidizedto N-(phosphonomethyl)glycine product, thereby producing an oxidationreaction mixture comprising the N-(phosphonomethyl)glycine product. Thevolumetric ratio of the liquid phase holdup in the fixed bed to thetotal bed volume is between about 0.1 and about 0.5.

In another embodiment, the continuous process comprises introducing anaqueous feed stream comprising the N-(phosphonomethyl)iminodiacetic acidsubstrate into an oxidation reaction zone comprising a fixed bedcontaining an oxidation catalyst. An O₂-containing gas is introducedinto the oxidation reaction zone wherein theN-(phosphonomethyl)iminodiacetic acid substrate is continuously oxidizedto the N-(phosphonomethyl)glycine product, thereby producing anoxidation reaction mixture comprising the N-(phosphonomethyl)glycineproduct. The partial pressure of oxygen at the liquid exit of the fixedbed is not greater than about 100 psia.

In a further embodiment, the continuous process comprises introducing anaqueous feed stream comprising the N-(phosphonomethyl)iminodiacetic acidsubstrate into an oxidation reaction zone comprising a fixed bedcontaining an oxidation catalyst. An O₂-containing gas is introducedinto the oxidation reaction zone wherein theN-(phosphonomethyl)iminodiacetic acid substrate is continuously oxidizedto the N-(phosphonomethyl)glycine product, thereby producing anoxidation reaction mixture comprising the N-(phosphonomethyl)glycineproduct. The partial pressure of oxygen is not greater than about 50psia at any location in the fixed bed at which the concentration ofN-(phosphonomethyl)iminodiacetic acid substrate in the liquid phase islower than about 0.1 ppm.

In a further embodiment, the continuous process comprises introducing anaqueous feed stream comprising the N-(phosphonomethyl)iminodiacetic acidsubstrate into an oxidation reaction zone comprising a fixed bedcontaining an oxidation catalyst. The catalyst surface area to liquidholdup in the fixed bed is between about 100 and about 6000 m²/cm³. Anoxidizing agent is introduced into the oxidation reaction zone whereinthe N-(phosphonomethyl)iminodiacetic acid substrate is continuouslyoxidized to the N-(phosphonomethyl)glycine product, thereby producing anoxidation reaction mixture comprising the N-(phosphonomethyl)glycineproduct.

In a further embodiment, the continuous process comprises introducing anaqueous feed stream comprising the N-(phosphonomethyl)iminodiacetic acidsubstrate into an oxidation reaction zone comprising a fixed bedcontaining an oxidation catalyst. An O₂-containing gas is introducedinto the oxidation reaction zone wherein theN-(phosphonomethyl)iminodiacetic acid substrate is continuously oxidizedto the N-(phosphonomethyl)glycine product, thereby producing anoxidation reaction mixture comprising the N-(phosphonomethyl)glycineproduct. The integrated average partial pressure of oxygen along theliquid flow path in the fixed bed is at least about 50 psia and theintegrated average temperature of the liquid phase in the fixed bedbeing between about 80° C. and about 130° C.

In a still further embodiment, the continuous process comprisesintroducing an aqueous feed stream comprising theN-(phosphonomethyl)iminodiacetic acid substrate into an oxidationreaction zone comprising a fixed bed containing oxidation catalystbodies and other means for promoting gas/liquid mass transfer. AnO₂-containing gas is introduced into the oxidation reaction zone whereinthe N-(phosphonomethyl)iminodiacetic acid substrate is continuouslyoxidized to the N-(phosphonomethyl)glycine product, thereby producing anoxidation reaction mixture comprising the N-(phosphonomethyl)glycineproduct.

In yet a further embodiment, the continuous process comprisesintroducing a liquid phase feed stream comprising an aqueous feedmixture comprising the N-(phosphonomethyl)iminodiacetic acid substrateinto a primary oxidation reaction zone comprising a fixed bed containingan oxidation catalyst. An oxidizing agent is introduced into the primaryoxidation reaction zone wherein the N-(phosphonomethyl)iminodiaceticacid substrate is continuously oxidized to theN-(phosphonomethyl)glycine product, thereby producing a liquid phaseexit stream comprising a primary reaction mixture comprising theN-(phosphonomethyl)glycine product and unreactedN-(phosphonomethyl)iminodiacetic acid substrate. The liquid phase exitstream is withdrawn from the primary oxidation reaction zone. The rateof introduction of the liquid phase feed stream and withdrawal of theliquid phase exit stream is such that the liquid phase hourly spacevelocity in the fixed bed based on total bed volume is between about 0.5hr⁻¹ and about 20 hr⁻¹.

Other features of this invention will be in part apparent and in partpointed out hereinafter.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 shows an example of a cross-section of a honeycomb catalystsupport.

FIG. 2 is a schematic flow sheet of a continuous oxidation reactorsystem for oxidizing an N-(phosphonomethyl)iminodiacetic acid substrateto form an N-(phosphonomethyl)glycine product. The reactor systemcomprises a back-mixed oxidation reaction zone utilizing a heterogeneousparticulate catalyst slurry recycled in a loop independent from a heattransfer recirculation loop.

FIG. 2A is a schematic flow sheet of a continuous oxidation reactorsystem for oxidizing an N-(phosphonomethyl)iminodiacetic acid substrateto form an N-(phosphonomethyl)glycine product. The reactor systemcomprises a back-mixed oxidation reaction zone utilizing a heterogeneousparticulate catalyst slurry recycled in a loop independent from a heattransfer recirculation loop and including a flash tank and catalystrecycle tank.

FIG. 2B is a schematic flow sheet of a continuous oxidation reactorsystem for oxidizing an N-(phosphonomethyl)iminodiacetic acid substrateto form an N-(phosphonomethyl)glycine product. The reactor systemcomprises a back-mixed oxidation reaction zone utilizing a heterogeneousparticulate catalyst slurry recycled through a heat transferrecirculation loop.

FIG. 3 is a schematic flow sheet of a continuous oxidation reactorsystem for oxidizing an N-(phosphonomethyl)iminodiacetic acid substrateto form an N-(phosphonomethyl)glycine product. The reactor systemcomprises two back-mixed oxidation reaction zones staged in seriesutilizing a heterogeneous particulate catalyst slurry which flows fromthe first reaction zone to the second reaction zone and is recycled tothe first reaction zone.

FIG. 4 is a schematic flow sheet of a continuous oxidation reactorsystem for oxidizing an N-(phosphonomethyl)iminodiacetic acid substrateto form an N-(phosphonomethyl)glycine product. The reactor systemcomprises two back-mixed oxidation reaction zones staged in seriesutilizing a heterogeneous particulate catalyst slurry which flows fromthe first reaction zone to the second reaction zone and is recycled toboth reaction zones.

FIG. 5 is a schematic flow sheet of a continuous oxidation reactorsystem for oxidizing an N-(phosphonomethyl)iminodiacetic acid substrateto form an N-(phosphonomethyl)glycine product. The reactor systemcomprises two back-mixed oxidation reaction zones staged in seriesutilizing two independent heterogeneous particulate catalyst slurrymasses such that catalyst from the first reaction zone is recycled tothe first reaction zone and catalyst from the second reaction zone isrecycled to the second reaction zone.

FIG. 6 is a schematic flow sheet of a continuous oxidation reactorsystem for oxidizing an N-(phosphonomethyl)iminodiacetic acid substrateto form an N-(phosphonomethyl)glycine product. The reactor systemcomprises two back-mixed oxidation reaction zones staged in seriesutilizing a heterogeneous particulate catalyst slurry which is recycledfrom the first reaction zone to the first reaction zone and from thesecond reaction zone to both reaction zones.

FIG. 7 is a schematic of an ejector nozzle loop reactor which may beused in the continuous oxidation reactor system of the present inventionfor oxidizing an N-(phosphonomethyl)iminodiacetic acid substrate to forman N-(phosphonomethyl)glycine product.

FIG. 8 is a schematic of a fixed bed reactor which may be used in thecontinuous oxidation reactor system of the present invention foroxidizing an N-(phosphonomethyl)iminodiacetic acid substrate to form anN-(phosphonomethyl)glycine product.

FIG. 9 is a schematic of a circulating fluidized bed reactor which maybe used in the continuous oxidation reactor system of the presentinvention for oxidizing an N-(phosphonomethyl)iminodiacetic acidsubstrate to form an N-(phosphonomethyl)glycine product.

FIG. 10 is a schematic flow sheet of a continuous distributed reactorsystem for oxidizing an N-(phosphonomethyl)iminodiacetic acid substrateto form an N-(phosphonomethyl)glycine product. The reactor systemcomprises a plurality of reactors in which reacting mixture progressesin series from each reactor to the succeeding reactor in the series.

FIG. 11 is a schematic flow sheet of an integrated process for oxidizingan N-(phosphonomethyl)iminodiacetic acid substrate in a reactor systemto form an oxidation reaction mixture comprising anN-(phosphonomethyl)glycine product and for recovering theN-(phosphonomethyl)glycine product from the oxidation reaction mixtureusing a non-adiabatic heat-driven evaporative crystallizer.

FIG. 12 is a schematic flow sheet of an integrated process for oxidizingan N-(phosphonomethyl)iminodiacetic acid substrate in a reactor systemto form an oxidation reaction mixture comprising anN-(phosphonomethyl)glycine product and for recovering theN-(phosphonomethyl)glycine product from the oxidation reaction mixtureusing an adiabatic crystallizer.

FIG. 12A is a schematic flow sheet of an adiabatic crystallizer systemused to recover an N-(phosphonomethyl)glycine product from an oxidationreaction mixture.

FIG. 13 is a schematic flow sheet of an integrated process for oxidizingan N-(phosphonomethyl)iminodiacetic acid substrate in a reactor systemto form an oxidation reaction mixture comprising anN-(phosphonomethyl)glycine product and for recovering theN-(phosphonomethyl)glycine product from the oxidation reaction mixtureusing a combination of an adiabatic crystallizer and a non-adiabaticheat-driven evaporative crystallizer operated in series.

FIG. 14 is a schematic flow sheet of an integrated process for oxidizingan N-(phosphonomethyl)iminodiacetic acid substrate in a reactor systemto form an oxidation reaction mixture comprising anN-(phosphonomethyl)glycine product and for recovering theN-(phosphonomethyl)glycine product from the oxidation reaction mixtureusing a combination of an adiabatic crystallizer and a non-adiabaticheat-driven evaporative crystallizer operated in semi-parallel.

FIG. 14A is a schematic flow sheet of an integrated process foroxidizing an N-(phosphonomethyl)iminodiacetic acid substrate to form anN-(phosphonomethyl)glycine product and for recovering theN-(phosphonomethyl)glycine product using a combination of an adiabaticcrystallizer and a non-adiabatic heat-driven evaporative crystallizeroperated in semi-parallel. The N-(phosphonomethyl)iminodiacetic acidsubstrate is oxidized in a primary reactor system to form an oxidationreaction mixture comprising the N-(phosphonomethyl)glycine product andunreacted N-(phosphonomethyl)iminodiacetic acid substrate. A primaryfraction of the oxidation reaction mixture from the primary reactorsystem is introduced into the adiabatic crystallizer, while unreactedN-(phosphonomethyl)iminodiacetic acid substrate in a secondary oxidationreactor feed fraction of the oxidation reaction mixture is oxidized in asecondary reactor system to form additional N-(phosphonomethyl)glycineproduct before being passed to the non-adiabatic crystallizer.

FIG. 15 shows the effect on the formic acid by-product concentrationprofile over 20 batch reaction runs caused by a one-time introduction ofbismuth oxide directly into an N-(phosphonomethyl)iminodiacetic acidoxidation reaction mixture. Here, the catalyst concentration in thereaction mixture was 0.5% by weight, and the catalyst contained 5% byweight platinum and 0.5% by weight iron.

FIG. 16 shows the effect on the formic acid by-product concentrationprofile over 30 batch reaction runs caused by a one-time introduction ofbismuth oxide directly into an N-(phosphonomethyl)iminodiacetic acidoxidation reaction mixture. Here, the catalyst concentration in thereaction mixture was 0.75% by weight, and the catalyst contained 5% byweight platinum and 1% by weight tin.

FIG. 17 shows the effect on the formaldehyde by-product concentrationprofile over 30 batch reaction runs caused by a one-time introduction ofbismuth oxide directly into an N-(phosphonomethyl)iminodiacetic acidoxidation reaction mixture. Here, the catalyst concentration in thereaction mixture was 0.75% by weight, and the catalyst contained 5% byweight platinum and 1% by weight tin.

FIG. 18 shows the effect on the N-methyl-N-(phosphonomethyl)glycine(NMG) by-product concentration profile over 30 batch reaction runscaused by a one-time introduction of bismuth oxide directly into anN-(phosphonomethyl)iminodiacetic acid oxidation reaction mixture. Here,the catalyst concentration in the reaction mixture was 0.75% by weight,and the catalyst contained 5% by weight platinum and 1% by weight tin.

FIG. 19 shows the effect on formic acid, formaldehyde, andN-methyl-N-(phosphonomethyl)glycine (NMG) production during anN-(phosphonomethyl)iminodiacetic acid oxidation reaction caused bymixing bismuth oxide with an oxidation catalyst that had been used in133 previous batch N-(phosphonomethyl)iminodiacetic acid oxidationreaction runs. Here, the catalyst comprised 5% by weight platinum and0.5% by weight iron on a carbon support.

FIG. 20 shows the effect on formic acid, formaldehyde, andN-methyl-N-(phosphonomethyl)glycine (NMG) production during anN-(phosphonomethyl)iminodiacetic acid oxidation reaction caused bymixing bismuth oxide with an oxidation catalyst that had been used in 30previous batch N-(phosphonomethyl)iminodiacetic acid oxidation reactionruns. Here, the catalyst comprised 5% by weight platinum and 1% byweight tin on a carbon support.

FIG. 21 shows the effect on the formic acid by-product concentrationprofile over 107 batch reaction runs caused by a one-time mixing ofbismuth oxide with a catalyst containing 5% by weight platinum and 1% byweight tin.

FIG. 22 shows the effect on the formaldehyde by-product concentrationprofile over 107 batch reaction runs caused by a one-time mixing ofbismuth oxide with a catalyst containing 5% by weight platinum and 1% byweight tin.

FIG. 23 shows the effect on the N-methyl-N-(phosphonomethyl)glycine(NMG) by-product concentration profile over 107 reaction runs caused bya one-time mixing of bismuth oxide with a catalyst containing 5% byweight platinum and 1% by weight tin.

FIG. 24 shows profiles of formaldehyde and formic acid in the productliquid of Example 21.

FIG. 25 shows profiles of glyphosate andN-(phosphonomethyl)iminodiacetic acid in the product liquid of Example22.

FIG. 26 shows profiles of glyphosate andN-(phosphonomethyl)iminodiacetic acid in the product liquid of Example23.

FIG. 27 is a block flow diagram for the continuous reactor system usedin Example 24.

FIG. 28 is a block flow diagram for the continuous reactor system usedin Example 25.

FIG. 29 is a block flow diagram for the continuous reactor system usedin Example 28.

FIG. 30 is a block flow diagram for the continuous reactor system usedin Example 35.

FIG. 31 is a block flow diagram for the continuous reactor system usedin Example 36.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

In general, the processes of this invention comprise (1) oxidizing anN-(phosphonomethyl)iminodiacetic acid substrate in one or more oxidationreaction zones to form an N-(phosphonomethyl)glycine product, and/or (2)concentrating and/or purifying the N-(phosphonomethyl)glycine product.These steps, along with several other features of the more preferredembodiments, are outlined below.

The N-(phosphonomethyl)iminodiacetic acid substrate is oxidized byintroducing the substrate and an oxidizing agent (i.e., oxygen source)into a reactor system comprising one or more oxidation reaction zonescontaining an oxidation catalyst. The oxidation reaction proceedsgenerally in accordance wither the following equation:

wherein R¹, R², R³, and R⁴ are each independently hydrogen, anagronomically acceptable cation, hydrocarbyl, or substitutedhydrocarbyl.

A hydrocarbyl is any group consisting exclusively of carbon andhydrogen. The hydrocarbyl may be branched or unbranched, may besaturated or unsaturated, and may comprise one or more rings. Suitablehydrocarbyl groups include alkyl, alkenyl, alkynyl, and aryl groups.They also include alkyl, alkenyl, alkynyl, and aryl groups substitutedwith other aliphatic or cyclic hydrocarbyl groups, such as alkaryl,alkenaryl, and alkynaryl.

A substituted hydrocarbyl is any hydrocarbyl wherein at least onehydrogen atom has been substituted with (a) an atom other than hydrogen,or (b) a group of atoms containing at least one atom other thanhydrogen. For example, the hydrogen atom may be substituted with ahalogen atom, such as a chlorine or fluorine atom. The hydrogen atomalternatively may be substituted with an oxygen atom or a groupcontaining an oxygen atom to form, for example, a hydroxy group, anether, an ester, an anhydride, an aldehyde, a ketone, or a carboxylicacid. The hydrogen atom also may be replaced with a group containing anitrogen atom to form, for example, an amide or a nitro group. Inaddition, the hydrogen atom may be substituted with a group containing asulfur atom to form, for example, —SO₃H.

An agronomically acceptable cation is a cation that allowsagriculturally and economically useful herbicidal activity of anN-(phosphonomethyl)glycine anion. Such a cation may be, for example, analkali metal cation (e.g., a sodium or potassium ion), an ammonium ion,an isopropyl ammonium ion, a tetra-alkylammonium ion, a trialkylsulfonium ion, a protonated primary amine, a protonated secondary amine,or a protonated tertiary amine.

In a particularly preferred embodiment, R¹, R², R³, and R⁴ are eachindependently hydrogen or an agronomically acceptable cation, withhydrogen often being most preferred.

Various oxidizing agents may be used in accordance with this invention.These include, for example, peroxides (e.g., H₂O₂, benzoyl peroxide),hydroperoxides, peroxy acids, O₂-containing gases, and liquidscomprising dissolved oxygen. Typically, O₂-containing gases areespecially preferred. As used herein, an O₂-containing gas is anygaseous mixture comprising O₂ and optionally one or more diluents whichare non-reactive with the oxygen or with the substrate or product underthe reaction conditions. Examples of such gases are air; pure O₂; or O₂diluted with He, Ar, N₂, and/or other non-oxidizing gases. The oxygensource is most preferably an O₂-containing gas containing at least about95 mole % O₂, more preferably about at least about 98 mole % O₂, withthe remainder being one or more non-oxidizing gases (particularly N₂and/or Ar).

Preferred Oxidation Catalysts

A wide variety of oxidation catalysts may be used in accordance withthis invention. These include both homogeneous and heterogeneouscatalysts.

Various water-soluble tungsten salts, for example, can be used tocatalyze the oxidation of N-(phosphonomethyl)iminodiacetic acidsubstrates with H₂O₂. N-(phosphonomethyl)iminodiacetic acid also can beoxidized to a N-oxide intermediate with H₂O₂ in the presence of an acid(e.g., H₂SO₄) and heat. This N-oxide intermediate, in turn, can bedecomposed to form N-(phosphonomethyl)glycine in the presence of heatand various water-soluble ferrous, cuprous, tungsten, molybdenum, andvanadium salt catalysts. A general discussion related to the use of suchhomogeneous catalysts for the conversion ofN-(phosphonomethyl)iminodiacetic acid to N-(phosphonomethyl)glycine canbe found, for example, in Franz, et al., Glyphosate: A Unique GlobalHerbicide (ACS Monograph 189, 1997) at pp. 240-41.

It is typically more preferred to use a heterogeneous catalyst. Thispreference stems, at least in part, from the ease with which aheterogeneous catalyst can normally be separated from the reactionmixture following the oxidation. The literature is replete with suitableheterogeneous catalysts.

One of the first heterogeneous catalysts used for catalyzing theoxidative cleavage of N-(phosphonomethyl)iminodiacetic acid is disclosedby Franz in U.S. Pat. No. 3,950,402. Franz discloses thatN-(phosphonomethyl)glycine may be prepared by the liquid phase oxidativecleavage of N-(phosphonomethyl)iminodiacetic acid with O₂ in thepresence of a catalyst comprising a noble metal deposited on the surfaceof an activated carbon support.

Even though Franz's process generally produces an acceptable yield andpurity of N-(phosphonomethyl)glycine, it also suffers from a number ofproblems:

-   -   1. The costly noble metal in Franz's catalyst tends to be lost        into the reaction solution (i.e., leaching). This noble metal        leaching is the result of at least two factors: (a) under the        oxidation conditions of the reaction, some of the noble metal is        oxidized into a more soluble form; and (b) both the        N-(phosphonomethyl)iminodiacetic acid substrate and the        N-(phosphonomethyl)glycine product act as ligands that        solubilize the noble metal.    -   2. The N-(phosphonomethyl)glycine product often oxidizes to form        aminomethylphosphonic acid (AMPA), particularly as the        concentration of the N-(phosphonomethyl)iminodiacetic acid        substrate decreases. This obviously reduces yield of the desired        N-(phosphonomethyl)glycine product.

In U.S. Pat. No. 3,969,398, Hershman discloses that activated carbonalone, without the presence of a noble metal, may be used to effect theoxidative cleavage of N-(phosphonomethyl)iminodiacetic acid to formN-(phosphonomethyl)glycine. In U.S. Pat. No. 4,624,937, Chou furtherdiscloses that the activity of the carbon catalyst disclosed by Hershmanmay be increased by removing the oxides from the surface of the carboncatalyst before using it in the oxidation reaction. See also, U.S. Pat.No. 4,696,772 (providing a separate discussion by Chou regardingincreasing the activity of the carbon catalyst by removing oxides fromthe surface of the carbon catalyst). Although these processes obviouslydo not suffer from noble metal leaching, they do tend to produce greaterconcentrations of formaldehyde and formic acid by-product when used toeffect the oxidative cleavage of N-(phosphonomethyl)iminodiacetic acid.

Optimally, the formaldehyde and formic acid are simultaneously oxidizedto carbon dioxide and water as the N-(phosphonomethyl)iminodiacetic acidsubstrate is oxidized to N-(phosphonomethyl)glycine, thus giving thefollowing reaction:

Much attention has focused on catalysts comprising a noble metal on acarbon support for at least two reasons. With such catalysts, the carboncomponent primarily effects the oxidation ofN-(phosphonomethyl)iminodiacetic acid to form N-(phosphonomethyl)glycineand formaldehyde, while the noble metal component primarily effects theoxidation of formaldehyde and formic acid to form carbon dioxide andwater. The noble metal component also tends to reduce the rate ofdeactivation of the carbon. More specifically, when activated carbon isused alone, it tends to deactivate by as much as 10% per cycle or more.Without being bound by any particular theory, it is believed that thedeactivation of the activated carbon alone arises because the surface ofthe carbon support oxidizes under the reaction conditions. See Chou,U.S. Pat. No. 4,624,937. See also, Chou, U.S. Pat. No. 4,696,772(providing a separate discussion related to deactivation of activatedcarbon by oxidation of the surface of the carbon). In the presence ofthe noble metal, however, the rate of deactivation of the activatedcarbon is diminished. It is believed that the noble metal reacts withthe oxidant at a faster rate than the activated carbon surface, and,thus, preferentially removes the oxidant from solution before extensiveoxidation of the carbon surface can occur. Further, unlike many oxidespecies which form at activated carbon surfaces and require hightemperature treatments to be reduced, oxide species which form at thesurface of a noble metal typically are easily reduced by the reducingagents present in or added to the reaction mixture (e.g., the aminefragment cleaved, formaldehyde, formic acid, H₂, etc.), thus restoringthe noble metal surface to a reduced state. In this manner, the catalystadvantageously exhibits significantly longer life as long as the noblemetal is not lost by leaching or sintered (i.e., in the form ofundesirably thick layers or clumps) by processes such as dissolution andre-deposition or noble metal agglomeration.

Ramon et al. (U.S. Pat. No. 5,179,228) disclose an example of using anoble metal deposited on the surface of a carbon support. To reduce theproblem of leaching (which Ramon et al. report to be as great as 30%noble metal loss per cycle), Ramon et al. disclose flushing the reactionmixture with N₂ under pressure after the oxidation reaction is completedto cause re-deposition of the noble metal onto the surface of the carbonsupport. According to Ramon et al., N₂ flushing reduces the noble metalloss to less than 1%.

Felthouse (U.S. Pat. No. 4,582,650) discloses using 2 catalysts: (i) anactivated carbon to effect the oxidation ofN-(phosphonomethyl)iminodiacetic acid to N-(phosphonomethyl)glycine, and(ii) a co-catalyst to concurrently effect the oxidation of formaldehydeto carbon dioxide and water. The co-catalyst is an aluminosilicatesupport having a noble metal located within its pores. The pores aresized to exclude N-(phosphonomethyl)glycine and thereby prevent thenoble metal of the co-catalyst from being poisoned byN-(phosphonomethyl)glycine. According to Felthouse, use of these 2catalysts together allows for the simultaneous oxidation ofN-(phosphonomethyl)iminodiacetic acid to N-(phosphonomethyl)glycine andof formaldehyde to carbon dioxide and water.

Ebner et al., International Publication No. WO 99/43430, the entiredisclosure of which is incorporated herein by references, discloseoxidizing N-(phosphonomethyl)iminodiacetic acid substrates using adeeply reduced catalyst comprising a noble metal on a carbon support.Such a catalyst tends to exhibit improved resistance to noble metalleaching and increased destruction of undesirable by-products (e.g.,formaldehyde). The advantages of such catalysts make them particularlypreferred. Thus, much of the following discussion will focus on suchcatalysts. Nevertheless, it should be recognized that the features ofthis invention may be generally applied using the wide variety ofhomogeneous and heterogeneous catalysts described above.

Oxygen-containing functional groups (e.g., carboxylic acids, ethers,alcohols, aldehydes, lactones, ketones, esters, amine oxides, andamides) at the surface of the carbon support tend to increase noblemetal leaching and potentially increase noble metal sintering duringliquid phase oxidation reactions, and, thus, reduce the ability of thecatalyst to oxidize oxidizable substrates, particularly formaldehyde andformic acid during the N-(phosphonomethyl)iminodiacetic acid oxidationreaction. As used herein, an oxygen-containing functional group is atthe surface of the carbon support if it is bound to an atom of thecarbon support and is able to chemically or physically interact withcompositions within the reaction mixture or with the metal atomsdeposited on the carbon support.

Many of the oxygen-containing functional groups that reduce noble metalresistance to leaching and sintering and reduce the activity of thecatalyst desorb from the carbon support as carbon monoxide when thecatalyst is heated at a high temperature (e.g., 900° C.) in an inertatmosphere (e.g., helium or argon). Thus, measuring the amount of COdesorption from a fresh catalyst (i.e., a catalyst that has notpreviously been used in a liquid phase oxidation reaction) under hightemperatures is one method that may be used to analyze the surface ofthe catalyst to predict noble metal retention and maintenance ofcatalyst activity. One way to measure CO desorption is by usingthermogravimetric analysis with in-line mass spectroscopy (TGA-MS).Preferably, no greater than about 1.2 mmole of carbon monoxide per gramof catalyst desorb from the catalyst when a dry, fresh sample of thecatalyst in a helium atmosphere is subjected to a temperature which isincreased from 20 to 900° C. at 10° C. per minute, and then heldconstant at 900° C. for 30 minutes. More preferably, no greater thanabout 0.7 mmole of carbon monoxide per gram of fresh catalyst desorbunder those conditions, even more preferably no greater than about 0.5mmole of carbon monoxide per gram of fresh catalyst desorb, and mostpreferably no greater than about 0.3 mmole of carbon monoxide per gramof fresh catalyst desorb. A catalyst is considered dry when the catalysthas a moisture content of less than 1% by weight. Typically, a catalystmay be dried by placing it into a N₂ purged vacuum of 25 inches of Hgand a temperature of 120° C. for 16 hours.

Measuring the number of oxygen atoms at the surface of a fresh catalystsupport is another method which may be used to analyze the catalyst topredict noble metal retention and maintenance of catalytic activity.Using, for example, x-ray photoelectron spectroscopy, a surface layer ofthe support which is about 50 Å in thickness is analyzed. Presentlyavailable equipment used for x-ray photoelectron spectroscopy typicallyis accurate to within ±20%. Typically, a ratio of carbon atoms to oxygenatoms at the surface (as measured by presently available equipment forx-ray photoelectron spectroscopy) of at least about 20:1 (carbonatoms:oxygen atoms) is suitable. Preferably, however, the ratio is atleast about 30:1, more preferably at least about 40:1, even morepreferably at least about 50:1, and most preferably at least about 60:1.In addition, the ratio of oxygen atoms to metal atoms at the surface(again, as measured by presently available equipment for x-rayphotoelectron spectroscopy) preferably is less than about 8:1 (oxygenatoms:metal atoms). More preferably, the ratio is less than 7:1, evenmore preferably less than about 6:1, and most preferably less than about5:1.

In general, the carbon supports used in the present invention are wellknown in the art. Activated, non-graphitized carbon supports arepreferred. These supports are characterized by high adsorptive capacityfor gases, vapors, and colloidal solids and relatively high specificsurface areas. The support suitably may be a carbon, char, or charcoalproduced by means known in the art, for example, by destructivedistillation of wood, peat, lignite, coal, nut shells, bones, vegetable,or other natural or synthetic carbonaceous matter, but preferably isactivated to develop adsorptive power. Activation usually is achieved byheating to high temperatures (800-900° C.) with steam or with carbondioxide which brings about a porous particle structure and increasedspecific surface area. In some cases, hygroscopic substances, such aszinc chloride and/or phosphoric acid or sodium sulfate, are added beforethe destructive distillation or activation, to increase adsorptivecapacity. Preferably, the carbon content of the carbon support rangesfrom about 10% for bone charcoal to about 98% for some wood chars andnearly 100% for activated carbons derived from organic polymers. Thenon-carbonaceous matter in commercially available activated carbonmaterials normally will vary depending on such factors as precursororigin, processing, and activation method. Many commercially availablecarbon supports contain small amounts of metals. Carbon supports havingthe fewest oxygen-containing functional groups at their surfaces aremost preferred.

The form of the support used in fixed bed reactors may varyconsiderably. For example, the carbon support may be in the form of amonolithic support. Suitable monolithic supports may have a wide varietyof shapes. A monolithic support may, for example, be in the form of areactor impeller. Even more preferably, such a support may also be, forexample, in the form of a screen or honeycomb having parallel channelsthrough which the feed mixture is passed. FIG. 1 shows an example of across-section of a honeycomb support. Although the cross-sections of thechannels in the honeycomb support of FIG. 1 are hexagonal in shape, ahoneycomb support as defined herein may alternatively (or additionally)comprise channels having other cross-section shapes (e.g., circular,oval, square, triangular rectangular, and the like). The channels of thehoneycomb support are preferably straight, and/or have a cross-sectionlarge enough so that they will not be clogged by a slurry containingsolid N-(phosphonomethyl)iminodiacetic acid substrate. Alternatively,the flow channels in a monolithic support may be irregular and without auniform flow direction (e.g., a random network of interconnecting flowchannels).

In a particularly preferred embodiment, the support is in the form ofparticulates. Because particulate supports are especially preferred,most of the following discussion focuses on embodiments which use aparticulate support. It should be recognized, however, that thisinvention is not limited to the use of particulate supports.

Suitable particulate supports may have a wide variety of shapes. Forexample, such supports may be in the form of pellets, granules andpowders. Pellet supports typically have a particle size of from about 1mm to about 10 mm. Preferably, the support is in the form of a powder.These particulate supports may be used in a reactor system as freeparticles, or, alternatively, may be bound to a structure in the reactorsystem, such as a screen or an impeller.

Typically, a support which is in particulate form comprises a broad sizedistribution of particles. For powders, preferably at least about 95% ofthe particles are from about 2 to about 300 μm in their largestdimension, more preferably at least about 98% of the particles are fromabout 2 to about 200 μm in their largest dimension, and most preferablyabout 99% of the particles are from about 2 to about 150 μm in theirlargest dimension with about 95% of the particles being from about 3 toabout 100 μm in their largest dimension. Particles greater than about200 μm in their largest dimension tend to fracture into super-fineparticles (i.e., less than 2 μm in their largest dimension), which aredifficult to recover.

The specific surface area of the carbon support, measured by the BET(Brunauer-Emmett-Teller) method using N₂, is preferably from about 10 toabout 3,000 m²/g (surface area of carbon support per gram of carbonsupport), more preferably from about 500 to about 2,100 m²/g, and stillmore preferably from about 750 to about 2,100 m²/g. In some embodiments,the most preferred specific area is from about 750 to about 1,750 m²/g.

The pore volume of the support may vary widely. Using the measurementmethod described in Example 1, the pore volume preferably is from about0.1 to about 2.5 ml/g (pore volume per gram of catalyst), morepreferably from about 0.2 to about 2.0 ml/g, and most preferably fromabout 0.4 to about 1.7 ml/g. Catalysts comprising supports with porevolumes greater than about 2.5 ml/g tend to fracture easily. On theother hand, catalysts comprising supports having pore volumes less than0.1 ml/g tend to have small surface areas and therefore low activity.

Carbon supports for use in the present invention are commerciallyavailable from a number of sources. The following is a listing of someof the activated carbons which may be used with this invention: DarcoG-60 Spec and Darco X (ICI-America, Wilmington, Del.); Norit SG Extra,Norit EN4, Norit EXW, Norit A, Norit Ultra-C, Norit ACX, and Norit 4×14mesh (Amer. Norit Co., Inc., Jacksonville, Fla.); Gl-9615, VG-8408,VG-8590, NB-9377, XZ, NV, and JV (Bamebey-Cheney, Columbus, Ohio); BLPulv., PWA Pulv., Calgon C 450, and PCB Fines (Pittsburgh ActivatedCarbon, Div. of Calgon Corporation, Pittsburgh, Pa.); P-100 (No. Amer.Carbon, Inc., Columbus, Ohio); Nuchar Conn., Nuchar C-1000 N, NucharC-190 A, Nuchar C-115 A, and Nuchar SA-30 (Westvaco Corp., CarbonDepartment, Covington, Va.); Code 1551 (Baker and Adamson, Division ofAllied Amer. Norit Co., Inc., Jacksonville, Fla.); Grade 235, Grade 337,Grade 517, and Grade 256 (Witco Chemical Corp., Activated Carbon Div.,New York, N.Y.); and Columbia SXAC (Union Carbide New York, N.Y.).

The carbon support preferably has one or more noble metal(s) at itssurface. Preferably, the noble metal(s) is selected from the groupconsisting of platinum (Pt), palladium (Pd), ruthenium (Ru), rhodium(Rh), iridium (Ir), silver (Ag), osmium (Os), and gold (Au). In general,platinum and palladium are more preferred, and platinum is mostpreferred. Because platinum is presently the most preferred noble metal,the following discussion will be directed primarily to embodiments usingplatinum. It should be understood, however, that the same discussion isgenerally applicable to the other noble metals and combinations thereof.It also should be understood that the term noble metal as used hereinmeans the noble metal in its elemental state as well as the noble metalin any of its various oxidation states.

The concentration of the noble metal deposited at the surface of thecarbon support may vary within wide limits. Preferably, it is in therange of from about 0.5 to about 20 wt. % ([mass of noble metal÷totalmass of catalyst]×100%), more preferably from about 2.5 to about 10 wt.%, and most preferably from about 3 to about 7.5 wt. %. Ifconcentrations less than 0.5 wt. % are used during theN-(phosphonomethyl)iminodiacetic acid oxidation reaction, there tends tobe less formaldehyde oxidized, and therefore a greater amount ofN-methyl-N-(phosphonomethyl)glycine produced, thereby reducing theN-(phosphonomethyl)glycine yield. On the other hand, at concentrationsgreater than about 20 wt. %, layers and clumps of noble metal tend toform. Thus, there are fewer surface noble metal atoms per total amountof noble metal used. This tends to reduce the activity of the catalystand is an uneconomical use of the costly noble metal.

The dispersion of the noble metal at the surface of the carbon supportpreferably is such that the concentration of surface noble metal atomsis from about 10 to about 400 μmole/g (μmole of surface noble metalatoms per gram of catalyst), more preferably, from about 10 to about 150μmole/g, and most preferably from about 15 to about 100 μmole/g. Thismay be determined, for example, by measuring chemisorption of H₂ or COusing a Micromeritics ASAP 2010C (Micromeritics, Norcross, Ga.) or anAltamira AMI100 (Zeton Altamira, Pittsburgh, Pa.).

Preferably, the noble metal is at the surface of the carbon support inthe form of metal particles. At least about 90% (number density) of thenoble metal particles at the surface of the carbon support arepreferably from about 0.5 to about 35 nm in their largest dimension,more preferably from about 1 to about 20 nm in their largest dimension,and most preferably from about 1.5 to about 10 nm in their largestdimension. In a particularly preferred embodiment, at least about 80% ofthe noble metal particles at the surface of the carbon support are fromabout 1 to about 15 nm in their largest dimension, more preferably fromabout 1.5 to about 10 nm in their largest dimension, and most preferablyfrom about 1.5 to about 7 nm in their largest dimension. If the noblemetal particles are too small, there tends to be an increased amount ofleaching when the catalyst is used in an environment that tends tosolubilize noble metals, as is the case when oxidizingN-(phosphonomethyl)iminodiacetic acid to formN-(phosphonomethyl)glycine. On the other hand, as the particle sizeincreases, there tends to be fewer noble metal surface atoms per totalamount of noble metal used. As discussed above, this tends to reduce theactivity of the catalyst and is also an uneconomical use of the costlynoble metal.

In addition to the noble metal, at least one promoter may be at thesurface of the carbon support. As defined herein, a promoter is a metalthat tends to increase catalyst selectivity, activity, and/or stability.A promoter additionally may reduce noble metal leaching. Although thepromoter usually is deposited onto the surface of the carbon support ina promoter deposition step, the carbon support itself may also (oralternatively) naturally contain a promoter. A promoter which isdeposited or exists naturally on the catalyst surface before the carbonsupport surface is finally reduced (described below) is referred toherein as a catalyst-surface promoter.

The catalyst-surface promoter may, for example, be an additional noblemetal(s) at the surface of the carbon support. For example, depending onthe application, ruthenium and palladium may act as catalyst-surfacepromoters on a catalyst comprising platinum deposited at a carbonsupport surface. The catalyst-surface promoter(s) alternatively may be,for example, a metal selected from the group consisting of tin (Sn),cadmium (Cd), magnesium (Mg), manganese (Mn), nickel (Ni), aluminum(Al), cobalt (Co), bismuth (Bi), lead (Pb), titanium (Ti), antimony(Sb), selenium (Se), iron (Fe), rhenium (Re), zinc (Zn), cerium (Ce),zirconium (Zr), tellurium (Te), and germanium (Ge). Preferably, thecatalyst-surface promoter is selected from the group consisting ofbismuth, iron, tin, tellurium and titanium. In a particularly preferredembodiment, the catalyst-surface promoter is tin. In anotherparticularly preferred embodiment, the catalyst-surface promoter isiron. In an additional preferred embodiment, the catalyst-surfacepromoter is titanium. In a further particularly preferred embodiment,the catalyst comprises both iron and tin at its surface. Use of iron,tin, or both generally (1) reduces noble metal leaching for a catalystused over several cycles, and (2) tends to increase and/or maintain theactivity of the catalyst when the catalyst is used to effect theoxidation of an N-(phosphonomethyl)iminodiacetic acid substrate.Catalysts comprising iron generally are most preferred because they tendto have the greatest activity and stability with respect to formaldehydeand formic acid oxidation.

In a preferred embodiment, the catalyst-surface promoter is more easilyoxidized than the noble metal (in instances where the catalyst-surfacepromoter is a noble metal as well, the catalyst-surface promoter noblemetal preferably is more easily oxidized than the non-promoter noblemetal). A promoter is more easily oxidized if it has a lower firstionization potential than the noble metal. First ionization potentialsfor the elements are widely known in the art and may be found, forexample, in the CRC Handbook of Chemistry and Physics (CRC Press, Inc.,Boca Raton, Fla.).

The amount of catalyst-surface promoter at the surface of the carbonsupport (whether associated with the carbon surface itself, metal, or acombination thereof) may vary within wide limits depending on, forexample, the noble metal(s) and catalyst-surface promoter(s) used.Typically, the weight percentage of the catalyst-surface promoter is atleast about 0.05% ([mass of catalyst-surface promoter÷total mass of thecatalyst]×100%). The weight percent of the catalyst-surface promoterpreferably is from about 0.05 to about 10%, more preferably from about0.1 to about 10%, still more preferably from about 0.1 to about 2%, andmost preferably from about 0.2 to about 1.5%. When the catalyst-surfacepromoter is tin, the weight percent most preferably is from about 0.5 toabout 1.5%. Catalyst-surface promoter weight percentages less than 0.05%generally do not promote the activity of the catalyst over an extendedperiod of time. On the other hand, weight percents greater than about10% tend to decrease the activity of the catalyst.

The molar ratio of noble metal to catalyst-surface promoter (and, ininstances where the catalyst-surface promoter is a noble metal as well,the molar ratio of the non-promoter noble metal to the catalyst-surfacepromoter noble metal) may also vary widely, depending on, for example,the noble metal(s) and catalyst-surface promoter(s) used. Preferably,the ratio is from about 1000:1 to about 0.01:1; more preferably fromabout 150:1 to about 0.05:1; still more preferably from about 50:1 toabout 0.05:1; and most preferably from about 10:1 to about 0.05:1. Forexample, a catalyst comprising platinum and iron preferably has a molarratio of platinum to iron of about 3:1.

In a particularly preferred embodiment of this invention, the noblemetal (e.g., Pt) is alloyed with at least one catalyst-surface promoter(e.g., Sn, Fe, or both) to form alloyed metal particles (and, ininstances where the catalyst-surface promoter is a noble metal as well,the non-promoter noble metal preferably is alloyed with thecatalyst-surface promoter noble metal). A catalyst comprising a noblemetal alloyed with at least one catalyst-surface promoter tends to haveall the advantages discussed above with respect to catalysts comprisinga catalyst-surface promoter in general. Catalysts comprising a noblemetal alloyed with at least one catalyst-surface promoter also tend toexhibit greater resistance to catalyst-surface promoter leaching andfurther stability from cycle to cycle with respect to formaldehyde andformic acid oxidation (See, e.g., Example 17).

The term alloy encompasses any metal particle comprising a noble metaland at least one catalyst-surface promoter, irrespective of the precisemanner in which the noble metal and catalyst-surface promoter atoms aredisposed within the particle (although it is generally preferable tohave a portion of the noble metal atoms at the surface of the alloyedmetal particle). The alloy may be, for example, any of the following:

-   -   1. An intermetallic compound. An intermetallic compound is        compound comprising a noble metal and a promoter (e.g., Pt₃Sn).    -   2. A substitutional alloy. A substitutional alloy has a single,        continuous phase, irrespective of the concentrations of the        noble metal and promoter atoms. Typically, a substitutional        alloy contains noble metal and promoter atoms which are similar        in size (e.g., platinum and silver; or platinum and palladium).        Substitutional alloys are also referred to as monophasic alloys.    -   3. A multiphasic alloy. A multiphasic alloy is an alloy that        contains at least two discrete phases. Such an alloy may        contain, for example Pt₃Sn in one phase, and tin dissolved in        platinum in a separate phase.    -   4. A segregated alloy. A segregated alloy is a metal particle        wherein the particle stoichiometry varies with distance from the        surface of the metal particle.    -   5. An interstitial alloy. An interstitial alloy is a metal        particle wherein the noble metal and promoter atoms are combined        with non-metal atoms, such as boron, carbon, silicon, nitrogen,        phosphorus, etc.

Preferably, at least about 80% (number density) of the alloyed metalparticles are from about 0.5 to about 35 nm in their largest dimension,more preferably from about 1 to about 20 nm in their largest dimension,still more preferably from about 1 to about 15 nm in their largestdimension, and most preferably from about 1.5 to about 7 nm in theirlargest dimension.

The alloyed metal particles need not have a uniform composition; thecompositions may vary from particle to particle, or even within theparticles themselves. In addition, the catalyst may further compriseparticles consisting of the noble metal alone or the catalyst-surfacepromoter alone. Nevertheless, it is preferred that the composition ofmetal particles be substantially uniform from particle to particle andwithin each particle, and that the number of noble metal atoms inintimate contact with catalyst-surface promoter atoms be maximized. Itis also preferred, although not essential, that the majority of noblemetal atoms be alloyed with a catalyst-surface promoter, and morepreferred that substantially all of the noble metal atoms be alloyedwith a catalyst-surface promoter. It is further preferred, although notessential, that the alloyed metal particles be uniformly distributed atthe surface of the carbon support.

Regardless of whether the catalyst-surface promoter is alloyed to thenoble metal, it is presently believed that the catalyst-surface promotertends to become oxidized if the catalyst is exposed to an oxidant over aperiod of time. For example, an elemental tin catalyst-surface promotertends to oxidize to form Sn(II)O, and Sn(II)O tends to oxidize to formSn(IV)O₂. This oxidation may occur, for example, if the catalyst isexposed to air for more than about 1 hour. Although suchcatalyst-surface promoter oxidation has not been observed to have asignificant detrimental effect on noble metal leaching, noble metalsintering, catalyst activity, or catalyst stability, it does makeanalyzing the concentration of detrimental oxygen-containing functionalgroups at the surface of the carbon support more difficult. For example,as discussed above, the concentration of detrimental oxygen-containingfunctional groups (i.e., oxygen-containing functional groups that reducenoble metal resistance to leaching and sintering, and reduce theactivity of the catalyst) may be determined by measuring (using, forexample, TGA-MS) the amount of CO that desorbs from the catalyst underhigh temperatures in an inert atmosphere. However, it is presentlybelieved that when an oxidized catalyst-surface promoter is present atthe surface, the oxygen atoms from the oxidized catalyst-surfacepromoter tend to react with carbon atoms of the support at hightemperatures in an inert atmosphere to produce CO, thereby creating theillusion of more detrimental oxygen-containing functional groups at thesurface of the support than actually exist. Such oxygen atoms of anoxidized catalyst-surface promoter also can interfere with obtaining areliable prediction of noble metal leaching, noble metal sintering, andcatalyst activity from the simple measurement (via, for example, x-rayphotoelectron spectroscopy) of oxygen atoms at the catalyst surface.

Thus, when the catalyst comprises at least one catalyst-surface promoterwhich has been exposed to an oxidant and thereby has been oxidized(e.g., when the catalyst has been exposed to air for more than about 1hour), it is preferred that the catalyst-surface promoter first besubstantially reduced (thereby removing the oxygen atoms of the oxidizedcatalyst-surface promoter from the surface of the catalyst) beforeattempting to measure the amount of detrimental oxygen-containingfunctional groups at the surface of the carbon support. This reductionpreferably is achieved by heating the catalyst to a temperature of 500°C. for 1 hour in an atmosphere consisting essentially of H₂. Themeasurement of detrimental oxygen-containing functional groups at thesurface preferably is performed (a) after this reduction, and (b) beforethe surface is exposed to an oxidant following the reduction. Mostpreferably, the measurement is taken immediately after the reduction.

The preferred concentration of metal particles at the surface of thecarbon support depends, for example, on the size of the metal particles,the specific surface area of the carbon support, and the concentrationof noble metal on the catalyst. It is presently believed that, ingeneral, the preferred concentration of metal particles is roughly fromabout 3 to about 1,500 particles/μm² (i.e., number of metal particlesper μm² of surface of carbon support), particularly where: (a) at leastabout 80% (number density) of the metal particles are from about 1.5 toabout 7 nm in their largest dimension, (b) the carbon support has aspecific surface area of from about 750 to about 2100 m²/g (i.e., m² ofsurface of carbon support per gram of carbon support), and (c) theconcentration of noble metal at the carbon support surface is from about1 to about 10 wt. % ([mass of noble metal÷total mass of catalyst]×100%).In more preferred embodiments, narrower ranges of metal particleconcentrations and noble metal concentrations are desired. In one suchembodiment, the concentration of metal particles is from about 15 toabout 800 particles/μm², and the concentration of noble metal at thecarbon support surface is from about 2 to about 10 wt. %. In an evenmore preferred embodiment, the concentration of metal particles is fromabout 15 to about 600 particles/μm², and the concentration of noblemetal at the carbon support surface is from about 2 to about 7.5 wt. %.In the most preferred embodiment, the concentration of the metalparticles is from about 15 to about 400 particles/μm², and theconcentration of noble metal at the carbon support surface is about 5wt. %. The concentration of metal particles at the surface of the carbonsupport may be measured using methods known in the art.

The surface of the carbon support preferably is deoxygenated before thenoble metal is deposited onto it. Preferably, the surface isdeoxygenated using a high-temperature deoxygenation treatment. Such atreatment may be a single-step or a multi-step scheme which, in eithercase, results in an overall chemical reduction of oxygen-containingfunctional groups at the surface of the carbon support.

In a two-step high-temperature deoxygenation treatment, the carbonsupport preferably is first treated with a gaseous or liquid phaseoxidizing agent to convert oxygen-containing functionalities inrelatively lower oxidation states (e.g., ketones, aldehydes, andalcohols) into functionalities in relatively higher oxidation states(e.g., carboxylic acids), which are easier to cleave from the surface ofthe catalyst at high temperatures. Representative liquid phase oxidizingagents include nitric acid, H₂O₂, chromic acid, and hypochlorite, withconcentrated nitric acid comprising from about 10 to about 80 grams ofHNO₃ per 100 grams of aqueous solution being preferred. Representativegaseous oxidants include molecular oxygen, ozone, nitrogen dioxide, andnitric acid vapors. Nitric acid vapors are the preferred oxidizingagent. With a liquid oxidant, temperatures of from about 60 to about 90°C. are appropriate, but with gaseous oxidants, it is often advantageousto use temperatures from about 50 to about 500° C. or even greater. Thetime during which the carbon is treated with the oxidant can vary widelyfrom about 5 minutes to about 10 hours. Preferably, the reaction time isfrom about 30 minutes to about 6 hours. Experimental results indicatethat carbon load, temperature, oxidant concentration, etc. in the firsttreatment step are not narrowly critical to achieving the desiredoxidation of the carbon material and thus may be governed by convenienceover a wide range. The highest possible carbon load is preferred foreconomic reasons.

In the second step, the oxidized carbon support is pyrolyzed (i.e.,heated) at a temperature preferably in the range of from about 500 toabout 1500° C., and more preferably from about 600 to about 1,200° C.,in a nitrogen, argon, helium, or other non-oxidizing environment (i.e.,an environment consisting essentially of no oxygen) to drive off theoxygen-containing functional groups from the carbon surface. Attemperatures greater than 500° C., an environment may be used whichcomprises a small amount of ammonia (or any other chemical entity whichwill generate NH₃ during pyrolysis), steam, or carbon dioxide which aidin the pyrolysis. As the temperature of the carbon support is cooled totemperatures less than 500° C., however, the presence ofoxygen-containing gases such as steam or carbon dioxide may lead to there-formation of surface oxides and thus, is preferably avoided.Accordingly, the pyrolysis is preferably conducted in a non-oxidizingatmosphere (e.g., nitrogen, argon, or helium). In one embodiment, thenon-oxidizing atmosphere comprises ammonia, which tends to produce amore active catalyst in a shorter time as compared to pyrolysis in theother atmospheres. The pyrolysis may be achieved, for example, using arotary kiln, a fluidized-bed reactor, or a conventional furnace.

The carbon support generally is pyrolyzed for a period of from about 5minutes to about 60 hours, preferably from about 10 minutes to about 6hours. Shorter times are preferred because prolonged exposure of thecarbon at elevated temperatures tends to reduce the activity of thecatalyst. Without being bound to any particular theory, it is presentlybelieved that prolonged heating at pyrolytic temperatures favors theformation of graphite, which is a less preferred form of a carbonsupport because it normally has less surface area. As discussed above, amore active catalyst typically may be produced in a shorter time byusing an atmosphere which comprises ammonia.

In a preferred embodiment of this invention, high-temperaturedeoxygenation is carried out in one step. This one-step treatment mayconsist of merely performing the pyrolysis step of the two-stephigh-temperature deoxygenation treatment discussed above. Morepreferably, however, the single-step treatment consists of pyrolyzingthe carbon support as described above while simultaneously passing a gasstream comprising N₂, NH₃ (or any other chemical entity which willgenerate NH₃ during pyrolysis), and steam over the carbon. Although itis not a critical feature of this invention, the flow rate of the gasstream preferably is fast enough to achieve adequate contact between thefresh gas reactants and the carbon surface, yet slow enough to preventexcess carbon weight loss and material waste. A non-reactive gas may beused as a diluent to prevent severe weight loss of the carbon.

Methods used to deposit the noble metal onto the surface of the carbonsupport are generally known in the art, and include liquid phase methodssuch as reaction deposition techniques (e.g., deposition via reductionof noble metal compounds, and deposition via hydrolysis of noble metalcompounds), ion exchange techniques, excess solution impregnation, andincipient wetness impregnation; vapor phase methods such as physicaldeposition and chemical deposition; precipitation; electrochemicaldeposition; and electroless deposition. See generally, Cameron, D. S.,Cooper, S. J., Dodgson, I. L., Harrison, B., and Jenkins, J. W. “Carbonsas Supports for Precious Metal Catalysts,” Catalysis Today, 7, 113-137(1990). Catalysts comprising noble metals at the surface of a carbonsupport also are commercially available, e.g., Aldrich Catalog No.20,593-1, 5% platinum on activated carbon (Aldrich Chemical Co., Inc.,Milwaukee, Wis.); Aldrich Catalog No. 20,568-0, 5% palladium onactivated carbon.

Preferably, the noble metal is deposited via a reactive depositiontechnique comprising contacting the carbon support with a solutioncomprising a salt of the noble metal, and then hydrolyzing the salt. Anexample of a suitable platinum salt which is relatively inexpensive ishexachloroplatinic acid (H₂PtCl₆). The use of this salt to depositplatinum onto a carbon support via hydrolytic deposition is illustratedin Example 3.

In one embodiment of this invention, the noble metal is deposited ontothe surface of the carbon support using a solution comprising a salt ofa noble metal in one of its more reduced oxidation states. For example,instead of using a salt of Pt(IV) (e.g., H₂PtCl₆), a salt of Pt(II) isused. In another embodiment, platinum in its elemental state (e.g.,colloidal platinum) is used. Using these more reduced metal precursorsleads to less oxidation of the carbon support and, therefore, lessoxygen-containing functional groups being formed at the surface of thesupport while the noble metal is being deposited onto the surface. Oneexample of a Pt(II) salt is K₂PtCl₄. Another potentially useful Pt(II)salt is diamminedinitrito platinum(II). Example 11 shows that using thissalt to deposit the noble metal produces a catalyst which is moreresistant to leaching than a catalyst prepared using H₂PtCl₆ as themetal precursor. Without being bound by any particular theory, it isbelieved that this is due to the fact that diamminedinitritoplatinum(II) generates ammonia in-situ during reduction which furtherpromotes removal of the oxygen-containing functional groups at thesurface of the carbon support. This benefit, however, should be weighedagainst a possible explosion danger associated with the use ofdiamminedinitrito platinum(II).

A catalyst-surface promoter(s) may be deposited onto the surface of thecarbon support before, simultaneously with, or after deposition of thenoble metal onto the surface. Methods used to deposit a promoter ontothe surface of the carbon support are generally known in the art, andinclude the same methods used to deposit a noble metal discussed above.In one embodiment, a salt solution comprising a promoter is used todeposit the catalyst-surface promoter. A suitable salt that may be usedto deposit bismuth is Bi(NO₃)₃.5H₂O, a suitable salt that may be used todeposit iron is FeCl₃.6H₂O, and a suitable salt that may be used todeposit tin is SnCl₂.2H₂O. It should be recognized that more than onecatalyst-surface promoter may be deposited onto the surface of thecarbon support. Examples 13, 14, 15, and 17 demonstrate depositing apromoter onto a carbon surface with a salt solution comprising apromoter. Example 18 demonstrates depositing more than one promoter(i.e., iron and tin) onto a carbon surface using salt solutionscomprising the promoters.

As noted above, a catalyst comprising a noble metal alloyed with atleast one catalyst-surface promoter is particularly preferred. There area variety of possible preparative techniques known in the art which maybe used to form a multi-metallic alloy at support surfaces. See, e.g.,V. Ponec & G. C. Bond, Catalysis by Metals and Alloys, “Studies inSurface Science and Catalysis,” Vol. 95 (B. Delmon. & J. T. Yates,advisory eds., Elsevier Science B.V., Amsterdam, Netherlands).

In one of the more preferred embodiments, reactive deposition is used toform metal particles containing a noble metal alloyed with acatalyst-surface promoter. Reactive deposition may comprise, forexample, reductive deposition wherein a surface of a carbon support iscontacted with a solution comprising: (a) a reducing agent; and (b) (i)a compound comprising the noble metal and a compound comprising thepromoter, or (ii) a compound comprising both the noble metal and thepromoter. A wide range of reducing agents may be used, such as sodiumborohydride, formaldehyde, formic acid, sodium formate, hydrazinehydrochloride, hydroxylamine, and hypophosphorous acid. Compoundscomprising a noble metal and/or a promoter include, for example:

-   -   1. Halide compounds. These include, for example, H₂PtCl₆,        K₂PtCl₄, Pt₂Br₆ ²⁻, K₂PdCl₄, AuCl₄ ¹⁻, RuCl₃, RhCl₃.3H₂O,        K₂RuCl₆, FeCl₃.6H₂O, (SnCl₃)¹⁻, SnCl₄, ReCl₆, FeCl₂, and TiCl₄.    -   2. Oxide and oxy chloride compounds. These include, for example,        RuO₄ ²⁻ and M₂SnO₄.    -   3. Nitrate compounds. These include, for example, Fe(NO₃)₃.    -   4. Amine complexes. These include, for example, [Pt(NH₃)₄]Cl₂,        [Pd(NH₃)₄]Cl₂, Pt(NH₃)₂Cl₂, Pt(NH₃)₄]PtCl₄, Pd(NH₂CH₂CH₂NH₂)Cl₂,        Pt(NH₂CH₂CH₂NH₂)₂Cl₂, and [Ru(NH₃)₅Cl]Cl₂.    -   5. Phosphine complexes. These include, for example,        Pt(P(CH₃)₃)₂Cl₂; IrClCO(P(C₆H₅)₃)₂; PtClH(PR₃)₂, wherein each R        is independently a hydrocarbyl, such as methyl, ethyl, propyl,        phenyl, etc    -   6. Organometallic complexes. These include, for example,        Pt₂(C₃H₆)₂Cl₄; Pd₂(C₂H₄)₂Cl₄; Pt(CH₃COO)₂, Pd(CH₃COO)₂;        K[Sn(HCOO)₃]; Fe(CO)₅; Fe₃(CO)₁₂; Fe₄(CO)₁₆; Sn₃(CH₃)₄; and        Ti(OR)₄, wherein each R is independently a hydrocarbyl, such as        methyl, ethyl, propyl, phenyl, etc.    -   7. Noble metal/promoter complexes. These include, for example,        Pt₃(SnCl₃)₂(C₈H₁₂)₃ and [Pt(SnCl₃)₅]³⁻.

In a particularly preferred embodiment, hydrolysis reactions are used todeposit a noble metal alloyed with a catalyst-surface promoter. In thisinstance, ligands containing the noble metal and promoter are formed,and then hydrolyzed to form well-mixed, metal oxide and metal hydroxideclusters at the surface of the carbon support. The ligands may beformed, for example, by contacting the surface of the support with asolution comprising (a) a compound comprising the noble metal and acompound comprising the promoter, or (b) a compound comprising both thenoble metal and the promoter. Suitable compounds comprising a noblemetal and/or a promoter are listed above with respect to reductivedeposition. Hydrolysis of the ligands may be achieved, for example, byheating (e.g., at a temperature of at least about 60° C.) the mixture.Example 17 further demonstrates the use of hydrolysis reactions todeposit a noble metal (i.e., platinum) alloyed with a catalyst-surfacepromoter (i.e., iron).

In addition to the above-described reactive deposition techniques, thereare many other techniques which may be used to form the alloy. Theseinclude, for example:

-   -   1. Forming the alloy by introducing metal compounds (which may        be simple or complex, and may be covalent or ionic) to the        surface of the support via impregnation, adsorption from a        solution, and/or ion exchange.    -   2. Forming the alloy by vacuum co-deposition of metal vapors        containing the noble metal and promoter onto the surface.    -   3. Forming the alloy by depositing one or metals onto a        pre-deposited metal belonging to Group 8, 9, or 10 of the        Periodic Table of the Elements (i.e., Fe, Co, Ni, Ru, Rh, Pd,        Os, Ir, and Pt) via, for example, electrolytic or electroless        plating.    -   4. Forming the alloy by: (a) depositing metal complexes        containing metals in the zero valence state (e.g., carbonyl,        pi-allyl, or cyclopentadienyl complexes of the noble metal and        of the promoter) at the surface of the carbon support; and (b)        removing the ligands by, for example, heating or reduction to        form the alloy particles at the surface.    -   5. Forming the alloy by contacting a solution containing a metal        compound (e.g., a metal chloride or a metal alkyl compound) with        a pre-deposited metal hydride containing a metal belonging to        Group 8, 9, or 10 of the Periodic Table of the Elements.    -   6. Forming the alloy by co-depositing, either simultaneously or        sequentially, metal complexes (either preformed or formed in        situ) containing the noble metal(s) and promoter(s) at the        surface of the carbon support.    -   7. Forming the alloy by pre-forming alloy particles as colloids        or aerosols, and then depositing the preformed alloy particles        at the surface of the carbon support. To illustrate, colloidal        particles containing platinum and iron may be easily formed by        boiling a dilute solution of H₂PtCl₆ and SnCl₂.2H₂O with a        sodium citrate solution. Protecting agents (e.g., carbohydrates,        polymers, lipophilic quaternary nitrogen salts) may be used to        effectively control metal alloy particle growth. This technique,        therefore, is often useful to form a narrow distribution of        alloy particle sizes.

It should be recognized that the above-discussed techniques for formingan alloy are simply illustrative, and not exhaustive. Using theteachings of this specification and the general knowledge of the art,one of ordinary skill in the art may routinely determine which of thenumerous alloy preparation techniques known in the art are suitable to aparticular use.

Regardless of the technique used to form the alloy, after the metalshave been deposited at the surface of the carbon support, it is oftenpreferable to dry the support using, for example, a sub-atmospheric,non-oxidizing environment (preferably, N₂, a noble gas, or both). Use ofa drying step is particularly preferred where the surface of the supportis to be subsequently reduced by heating the surface (and even morepreferred where the heating is to be conducted in a non-oxidizingenvironment). Preferably, the support is dried to reduce the moisturecontent of the support to less than about 5% by weight.

It should be recognized that reducing the surface of the carbon supportafter deposition of the noble metal(s) and catalyst-surface promoter(s)typically increases the extent of noble metal alloyed with acatalyst-surface promoter. Such reduction also often tends to increasethe number of particles falling within the preferred size range.

After the carbon support has been impregnated with the noble metal(s)(and catalyst-surface promoter(s), if any), the surface of the catalystpreferably is reduced. The surface of the catalyst suitably may bereduced, for example, by heating the surface at a temperature of atleast about 400° C. It is especially preferable to conduct this heatingin a non-oxidizing environment (e.g., nitrogen, argon, or helium). It isalso more preferred for the temperature to be greater than about 500° C.Still more preferably, the temperature is from about 550 to about 1,200°C., and most preferably from about 550 to about 900° C. Temperaturesless than 400° C. tend to be unsatisfactory for removing theoxygen-containing functional groups from the surface of the carbonsupport. On the other hand, temperatures greater than 1,200° C. tend toreduce the activity of the catalyst. Temperatures of from about 400 toabout 500° C. preferably are used only if the surface of the carbonsupport has a carbon atom to oxygen atom ratio of at least about 20:1before the noble metal is deposited onto the surface.

In a particularly preferred embodiment, the surface of the catalyst isreduced by a process comprising exposing the surface to a reducingenvironment. For example, before the heating, the catalyst sample may bepre-treated with a liquid phase reducing agent, such as formaldehyde orformic acid. Even more preferably, the heating is conducted in thepresence of a gas-phase reducing agent (the method of heating thecatalyst in the presence of a gas-phase reducing agent will sometimes bereferred to as high-temperature gas-phase reduction). Various gas-phasereducing agents may be used during the heating, including but notlimited to H₂, ammonia, and carbon monoxide. Hydrogen gas is mostpreferred because the small molecular size of hydrogen allows betterpenetration into the deepest pores of the carbon support. Preferably,the remainder of the gas consists essentially of a non-oxidizing gas,such as nitrogen, argon, or helium. The gas may comprise any finiteconcentration of H₂, although H₂ concentrations of less than 1.0% aredisadvantageous because of the time they tend to require to reduce thesurface of the support. Preferably, the gas comprises from about 5 toabout 50 volume % H₂, and most preferably from about 5 to about 25volume % H₂.

The preferred amount of time that the catalyst surface is heated dependson the mass transfer of the reducing agent to the catalyst surface. Whenthe reducing agent is a non-oxidizing gas comprising from about 10 toabout 20 volume % H₂, the surface preferably is heated for from about 15minutes to about 24 hours at from about 550 to about 900° C. with aspace velocity of from about 1 to about 5,000 hour⁻¹. More preferably,the space velocity is from about 10 to about 2,500 hour⁻¹, and even morepreferably from about 50 to about 750 hour⁻¹. In the most preferredembodiment, the heat treatment is conducted at the above preferredtemperatures and space velocities for from about 1 to about 10 hours⁻¹.Heating the surface at space velocities of less than 1 hour⁻¹ isdisadvantageous because the oxygen-containing functional groups at thesurface of the carbon support may not be sufficiently destroyed. On theother hand, heating the surface at space velocities greater than 5,000hour⁻¹ is uneconomical.

Pre-existing oxygen-containing functional groups at the surface of thecarbon support generally are not necessary, or even desired, to obtainadequate noble metal dispersion and retention. Without being bound byany particular theory, it is believed that this heating step enhancesthe platinum-carbon interaction on the catalyst by removingoxygen-containing functional groups at the surface of the carbonsupport, including those formed by depositing the noble metal onto thesurface. It is believed that these oxygen-containing functional groupsare unstable anchor sites for the noble metal because they tend tointerfere with the potentially stronger π interactions between the noblemetal and the carbon support. Heating alone will decompose and therebyremove many of the oxygen-containing functional groups at the surface ofthe carbon support. However, by heating the surface in the presence of areducing agent (e.g., H₂), more oxygen-containing functional groups areable to be eliminated.

If the carbon atom to oxygen atom ratio at the surface of the carbonsupport is less than about 20:1 before the noble metal is deposited ontothe surface of the support, the surface preferably is reduced using theabove-described high-temperature gas-phase reduction treatment at atemperature greater than 500° C., although the surface may optionally betreated with other reduction environments in addition tohigh-temperature gas-phase reduction. On the other hand, if the surfaceof the carbon support has a carbon atom to oxygen atom ratio which is atleast about 20:1 before the noble metal is deposited onto the surface,various alternative reduction environments may be used instead ofhigh-temperature gas-phase reduction.

The surface of the catalyst may be reduced, at least in part, bytreating it with an amine, such as urea, a solution comprising ammoniumions (e.g., ammonium formate or ammonium oxalate), or ammonia gas, withammonia gas or a solution comprising ammonium ions being most preferred.This amine treatment preferably is used in addition to other reductiontreatments, and most preferably is used before high-temperaturegas-phase reduction. In one such embodiment, the noble metal isdeposited onto the surface by treating it with a noble metal precursorsolution comprising ammonium ions. Alternatively, after the noble metalis deposited onto the surface of the support, the support may be washedwith a solution comprising ammonium ions or placed into contact with agas comprising ammonia. Most preferably, the catalyst surface is washedwith diluted aqueous ammonia after depositing the noble metal. In thisinstance, the catalyst is added to pure water and stirred for a fewhours to wet the surface of the catalyst. Next, while continuing to stirthe catalyst slurry, a solution comprising ammonium ions is added to thecatalyst slurry in an amount sufficient to produce a pH of greater than7, more preferably from about 8 to about 12, and most preferably fromabout 9.5 to about 11.0. Because the temperature and pressure are notcritical, this step preferably is performed at room temperature andatmospheric pressure. Example 10 further demonstrates this reductiontreatment.

Sodium borohydride (NaBH₄) also may be used to reduce the surface of thecatalyst. As with the amine treatment, this treatment preferably is usedin addition to other reduction treatments, and most preferably is usedbefore high-temperature gas-phase reduction. Preferably, afterdepositing the noble metal onto the surface of the support, the supportis washed with a solution of NaBH₄ in the presence of NaOH at a pH offrom about 8 to about 14 for about 15 to about 180 minutes. The amountof NaBH₄ used preferably is sufficient to reduce all the noble metal.Because the temperature and pressure are not critical, this steppreferably is performed at room temperature and atmospheric pressure.Example 12 further demonstrates this reduction treatment.

It should be recognized that any of the above treatments which may beused to reduce the surface of the catalyst also may be used todeoxygenate the surface of the carbon support before the noble metal isdeposited onto the surface.

In many processes, when it is desirable for a catalyst to contain apromoter, the promoter is pre-deposited onto the catalyst surface by,for example, the promoter deposition techniques described above (thisdeposition step is often performed by the manufacturer of the catalyst).This promoter deposition step, however, tends to add costs to thecatalyst preparation process. To avoid these additional costs, it hasbeen found that the benefits of a promoter (e.g., increased selectivity,activity, and/or catalyst stability) may be obtained by merely mixing apromoter (i.e., a supplemental promoter) directly with acarbon-supported, noble-metal-containing catalyst (particularly with thereduced catalysts described above). This mixing may, for example, beconducted directly in an oxidation reaction zone where theN-(phosphonomethyl)iminodiacetic acid substrate is oxidized.Alternatively, for example, this mixing may take place separately fromthe oxidation reaction, such as in a catalyst holding tank.

Particularly, it has been discovered that certain metals and/or metalcompounds function as supplemental promoters in the oxidation of anN-(phosphonomethyl)iminodiacetic acid substrate catalyzed by acarbon-supported, noble-metal-containing catalyst. It has been foundthat such supplemental promoters are effective in enhancing thecapability of noble metal on carbon catalysts for the oxidation of anN-(phosphonomethyl)iminodiacetic acid substrate to anN-(phosphonomethyl)glycine product wherein they are effective inenhancing catalysis of the desired conversion to N-(phosphonomethyl)glycine, the oxidation of by-product formaldehyde to formic acid, andthe oxidation of by-product formic acid to carbon dioxide.

Depending on the application, the supplemental promoter(s) may be, forexample, tin, cadmium, magnesium, manganese, ruthenium, nickel, copper,aluminum, cobalt, bismuth, lead, titanium, antimony, selenium, iron,rhenium, zinc, cerium, zirconium, tellurium, sodium, potassium,vanadium, gallium, tantalum, niobium, rubidium, cesium, lanthanum,and/or germanium. It is often more preferred for the supplementalpromoter(s) to be bismuth, lead, germanium, tellurium, titanium, copperand/or nickel.

In an especially preferred embodiment, the supplemental promoter isbismuth. It has been found in accordance with this invention that thepresence of bismuth is especially effective in enhancing the selectivityof a carbon-supported, noble-metal-containing catalyst (particularly thereduced catalyst described above) when it is used to catalyze theoxidation of an N-(phosphonomethyl)iminodiacetic acid substrate to forman N-(phosphonomethyl)glycine product. More specifically, it has beenfound that the presence of bismuth causes an increase in the amount offormic acid by-product that is catalytically oxidized. In some instances(particularly where the catalyst comprises tin as a catalyst-surfacepromoter), the presence of bismuth also has been found to cause anincrease in the amount of formaldehyde by-product that is catalyticallyoxidized. This increased destruction of one or both of theseby-products, in turn, causes less N-methyl-N-(phosphonomethyl)glycineby-product to be formed (it is believed that this stems from the factthat the formation of each molecule ofN-methyl-N(phosphonomethyl)glycine by-product requires either (a) twoformaldehyde molecules, or (b) a formic acid molecule and a formaldehydemolecule). Further, it has been found that in some instances(particularly where more than one supplemental promoter is used) thatthe presence of bismuth may also reduce the amount of noble metal thatleaches from the carbon support of the catalyst during the oxidation ofan N-(phosphonomethyl)iminodiacetic acid substrate.

In another preferred embodiment of this invention, tellurium is used asa supplemental promoter. As in the above embodiment incorporatingbismuth as a supplemental promoter, it has been found that the presenceof tellurium is also effective in enhancing the selectivity of acarbon-supported, noble-metal-containing catalyst (particularly thereduced catalyst described above) when it is used to catalyze theoxidation of an N-(phosphonomethyl)iminodiacetic acid substrate to forman N-(phosphonomethyl)glycine product. More particularly, it has beenfurther found that tellurium may increase the activity of the catalystin the oxidation of an N-(phosphonomethyl)iminodiacetic acid substrate.Further, it has also been found that noble metal leaching from thecarbon support of the catalyst may be reduced during the oxidation of anN-(phosphonomethyl)iminodiacetic acid substrate by the presence oftellurium in the reaction medium (particularly when bismuth is alsopresent).

In a most preferred embodiment, both bismuth and tellurium are used assupplemental promoters.

The mixing of the supplemental promoter and catalyst preferably isconducted in a liquid medium. As noted above, this mixing may, forexample, be conducted directly in a liquid reaction medium where theN-(phosphonomethyl)iminodiacetic acid substrate is being oxidized.Where, however, the oxidation reaction is carried out under pressure,the reaction vessel is normally sealed and it is consequently often morepreferred to mix the catalyst with the supplemental promoter separatelyfrom the reaction vessel, such as in a catalyst holding or recycle tank.

Typically, the supplemental promoter is introduced into the mixingliquid in the form of an inorganic or organic compound containing thesupplemental promoter. The promoter-containing compound may be solubleor insoluble in the liquid, but most typically is at least partiallysoluble. The functional group attached to the supplemental promoter atomis generally not critical (although it preferably is an agronomicallyacceptable functional group). Typically, for example, suitable compoundsinclude oxides, hydroxides, salts of inorganic hydracids, salts ofinorganic oxy-acids, salts of aliphatic or aromatic organic acids, andphenates.

Suitable bismuth-containing compounds, for example, include inorganic ororganic compounds wherein the bismuth atom(s) is at an oxidation levelgreater than 0 (e.g., 2, 3, 4 or 5), most preferably 3. Examples of suchsuitable bismuth compounds include:

-   -   1. Bismuth oxides. These include, for example, BiO, Bi₂O₃,        Bi₂O₄, Bi₂O₅, and the like.    -   2. Bismuth hydroxides. These include, for example, Bi(OH)₃ and        the like.    -   3. Bismuth salts of inorganic hydracids. These include, for        example, bismuth chloride (e.g., BiCl₃), bismuth bromide (e.g.,        BiBr₃), bismuth iodide (e.g., BiI₃), bismuth telluride (e.g.,        Bi₂Te₃), and the like. Bismuth halides are typically less        preferred because they tend to be corrosive to the process        equipment.    -   4. Bismuth salts of inorganic oxy-acids. These include, for        example, bismuth sulphite (e.g., Bi₂(SO₃)₃.Bi₂O₃.5H₂O), bismuth        sulphate (e.g., Bi₂(SO₄)₃), bismuthyl sulfate (e.g., (BiO)HSO₄),        bismuthyl nitrite (e.g., (BiO)NO₂.0.5H₂O), bismuth nitrate        (e.g., Bi(NO₃)₃.5H₂O, also known as bismuth nitrate        pentahydrate), bismuthyl nitrate (e.g., (BiO)NO₃, also known as        bismuth subnitrate, bismuth nitrate oxide, and bismuth        oxynitrate), double nitrate of bismuth and magnesium (e.g.,        2Bi(NO₃)₃.3Mg(NO₃)₂.24H₂O), bismuth phosphite (e.g.,        Bi₂(PO₃H)₃.3H₂O), bismuth phosphate (e.g., BiPO₄), bismuth        pyrophosphate (e.g., Bi₄(P₂O₇)₃), bismuthyl carbonate (e.g.,        (BiO)₂CO₃, also known as bismuth subcarbonate), bismuth        perchlorate (e.g., Bi(ClO₄)₃.5H₂O), bismuth antimonate (e.g.,        BiSbO₄), bismuth arsenate (e.g., Bi(AsO₄)₃), bismuth selenite        (e.g., Bi₂(SeO₂)₃), bismuth titanate (e.g., Bi₂O₃.2TiO₂), and        the like. These salts also include bismuth salts of oxy-acids        derived from transition metals, including, for example, bismuth        vanadate (e.g., BiVO₄), bismuth niobate (e.g., BiNbO₄), bismuth        tantalate (BiTaO₄), bismuth chromate (Bi₂(CrO₄), bismuthyl        dichromate (e.g., (BiO)₂Cr₂O₇), bismuthyl chromate (e.g.,        H(BiO)CrO₄), double chromate of bismuthyl and potassium (e.g.,        K(BiO)CrO₄), bismuth molybdate (e.g., Bi₂(MoO₄)₃), double        molybdate of bismuth and sodium (e.g., NaBi(MoO₄)₂), bismuth        tungstate (e.g., Bi₂(WO₄)₃), bismuth permanganate (e.g.,        Bi₂O₂(OH)MnO₄), bismuth zirconate (e.g., 2Bi₂O₃.3ZrO₂), and the        like.    -   5. Bismuth salts of aliphatic or aromatic organic acids. These        include, for example, bismuth acetate (e.g., Bi(C₂H₃O₂)₃),        bismuthyl propionate (e.g., (BiO)C₃H₅O₂), bismuth benzoate        (e.g., C₆H₅CO₂Bi(OH)₂), bismuthyl salicylate (e.g.,        C₆H₄CO₂(BiO)(OH)), bismuth oxalate (e.g., (C₂O₄)₃Bi₂), bismuth        tartrate (e.g., Bi₂(C₄H₄O₆)₃.6H₂O), bismuth lactate (e.g.,        (C₆H₉O₅)OBi.7H₂O), bismuth citrate (e.g., C₆H₅O₇Bi), and the        like.    -   6. Bismuth phenates. These include, for example, bismuth gallate        (e.g., C₇H₇O₇Bi), bismuth pyrogallate (e.g.,        C₆H₃(OH)₂(OBi)(OH)), and the like.    -   7. Miscellaneous other organic and inorganic bismuth compounds.        These include, for example, bismuth phosphide (e.g., BiP),        bismuth arsenide (Bi₃As₄), sodium bismuthate (e.g., NaBiO₃),        bismuth-thiocyanic acid (e.g., H₂(Bi(BNS)₅).H₃(Bi(CNS)₆)),        sodium salt of bismuth-thiocyanic acid, potassium salt of        bismuth-thiocyanic acid, trimethylbismuthine (e.g., Bi(CH₃)₃),        triphenylbismuthine (e.g., Bi(C₆H₅)₃), bismuth oxychloride        (e.g., BiOCl), bismuth oxyiodide (e.g., BiOI), and the like.

In a preferred embodiment, the bismuth compound is a bismuth oxide,bismuth hydroxide, or bismuth salt of an inorganic oxy-acid. Morepreferably, the bismuth compound is bismuth nitrate (e.g.,Bi(NO₃)₃.5H₂O), bismuthyl carbonate (e.g., (BiO)₂CO₃), or bismuth oxide(e.g., Bi₂O₃), with bismuth (III) oxide (i.e., Bi₂O₃) being mostpreferred because it contains no counterion which can contaminate thefinal reaction product.

Suitable tellurium-containing compounds, for example, include inorganicor organic compounds wherein the tellurium atom(s) is at an oxidationlevel greater than 0 (e.g., 2, 3, 4, 5 or 6), most preferably 4.Examples of such suitable tellurium compounds include:

-   -   1. Tellurium oxides. These include, for example, TeO₂, Te₂O₃,        Te₂O₅, TeO₃, and the like.    -   2. Tellurium salts of inorganic hydracids. These include, for        example, tellurium tetrachloride (e.g., TeCl₄), tellurium        tetrabromide (e.g., TeBr₄), tellurium tetraiodide (e.g., TeI₄),        and the like.    -   3. Tellurium salts of inorganic oxy-acids. These include, for        example, tellurious acid (e.g., H₂TeO₃), telluric acid (e.g.,        H₂TeO₄ or Te(OH)₆), tellurium nitrate (e.g., Te₂O₄.HNO₃), and        the like.    -   4. Miscellaneous other organic and inorganic tellurium        compounds. These include, for example, dimethyl tellurium        dichloride, lead tellurium oxide, tellurium isopropoxide,        ammonium tellurate, tellurium thiourea, and the like.

In a preferred embodiment, the tellurium compound is a tellurium oxideor tellurium salt of an inorganic hydracid. More preferably, thetellurium compound is tellurium dioxide (e.g., TeO₂), telluriumtetrachloride (e.g., TeCl₄), or telluric acid (e.g., Te(OH)₆), withtellurium tetrachloride being most preferred.

The preferred amount of the supplemental promoter introduced into thereaction zone depends on, for example, the mass of the carbon-supported,noble-metal-containing catalyst (i.e., the total mass of the carbonsupport, noble metal, and any other component of the catalyst); mass ofthe total reaction feed mixture; and the concentration of theN-(phosphonomethyl) iminodiacetic acid substrate.

In general, the ratio of the mass of the supplemental promoter to themass of the carbon-supported, noble-metal-containing catalyst charged tothe reactor(s) is preferably at least about 1:15000; more preferably atleast about 1:5000; even more preferably at least about 1:2500; and mostpreferably at least about 1:1000. Although it is feasible to practicethe present invention without detriment to the oxidation reaction whenratios of the mass of supplemental promoter to the mass of thecarbon-supported, noble-metal-containing catalyst are as great as about1:750, about 1:500, about 1:300, and even greater than about 1:50 or1:40, the preferred lower ratios described above have been found to beeffective for most applications, and particularly for the oxidation ofan N-(phosphonomethyl) iminodiacetic acid substrate.

The ratio of the mass of the supplemental promoter to the total reactionmass charged to the reactor is preferably at least about 1:1,000,000;more preferably at least about 1:100,000; even more preferably at leastabout 1:40,000; and most preferably from about 1:40,000 to about1:15,000. Although ratios greater than 1:8000 may normally be usedwithout detriment to the oxidation reaction, it is generally preferredfor the ratio to be less than 1:8000 (particularly where bismuth is thesupplemental promoter).

The ratio of the mass of the supplemental promoter to the mass of theN-(phosphonomethyl)iminodiacetic acid substrate charged to the reactoris preferably at least about 1:100,000; more preferably 1:10,000; evenmore preferably at least about 1:4000; and most preferably from about1:4000 to about 1:2000. Although ratios greater than 1:1000 may normallybe used without detriment to the oxidation reaction, it is generallypreferred for the ratio to be less than 1:1000 (particularly wherebismuth is the supplemental promoter).

Where a particulate noble metal on carbon catalyst is used for theoxidation of an N-(phosphonomethyl)iminodiacetic acid substrate, boththe catalyst and the supplemental promoter may be charged to an aqueousreaction medium containing the N-(phosphonomethyl)iminodiacetic acidsubstrate and oxygen. The supplemental promoter may be charged in a massratio to the catalyst charge of at least about 1:15,000, preferably atleast about 1:5000, more preferably at least about 1:2500, and mostpreferably at least about 1:1000. As oxidation of theN-(phosphonomethyl)iminodiacetic acid substrate proceeds, formaldehydeand formic acid by-products are generated. The catalyst is effective tocatalyze not only the oxidation of the N-(phosphonomethyl)iminodiaceticacid substrate but also the further oxidation of formaldehyde to formicacid, and formic acid to carbon dioxide. The presence of thesupplemental promoter is effective to enhance the catalytic oxidation ofthese by-products, especially for the conversion of formic acid to CO₂.

Where the oxidation reaction is conducted in a stirred tank reactor inwhich catalyst is slurried in the reaction medium, the catalyst isseparated from the reaction mixture, preferably by filtration, andrecycled to the reactor for further oxidation of theN-(phosphonomethyl)iminodiacetic acid substrate and the aforesaidby-products. Such a stirred tank reactor system may be operated ineither a batch or continuous mode. Alternatively, a fixed or fluidcatalyst bed can be used. In a continuous process, theN-(phosphonomethyl)iminodiacetic acid substrate, formaldehyde and formicacid are all oxidized in a continuous reaction zone to which an aqueousreaction medium comprising the N-(phosphonomethyl)imindiacetic acidsubstrate is continuously or intermittently supplied and a reactionmixture comprising an N-(phosphonomethyl)glycine product is continuouslyor intermittently withdrawn, the supplemental promoter beingcontinuously or intermittently introduced into the reaction zone.

It has been observed that addition of a discrete charge of supplementalpromoter to the first batch of series of successive batch reactioncycles is effective to enhance the activity of the catalyst foroxidation of formaldehyde and formic acid throughout the series ofreaction cycles, without further addition of supplemental promoter fromany external source. It has further been observed that the supplementalpromoter is present in the recycled catalyst, apparently having beendeposited thereon by adsorption to the noble metal and/or the carbonsupport. Only a fraction of the supplemental promoter added to the firstbatch of the series can be found on the catalyst after multiple cycles.However, when supplemental promoter is introduced into the first batchin the amounts described above, the fraction remaining on the catalystis apparently sufficient for promoting the oxidation of formaldehyde andformic acid throughout the series of batches in which the catalystrecycled from an earlier batch is substantially the sole source ofsupplemental promoter for the successive batch reaction cycles of theseries. It has been found that a single addition of supplementalpromoter in a mass ratio to the catalyst of approximately 1:2500 iseffective for promotion of by-product oxidation in series of 20 or more,typically 50 or more, more typically over 100, batch reaction cycles.Thereafter, a further discrete charge of supplemental promoteroptionally may be added to the reaction medium for a subsequent batchconstituting the first of another series of batch oxidation reactioncycles in which the recycle catalyst from an earlier batch of suchfurther series becomes substantially the sole source of promoter for thesuccessive batch reaction cycles of the further series of batchreactions.

Similarly, where supplemental promoter is added to the reaction mediumin a continuous stirred tank reactor, addition of supplemental promoterin a single discrete amount is effective to enhance the effectiveness ofthe catalyst for formaldehyde and formic acid oxidation throughoutmultiple reactor turnovers of a continuous reaction run. No furtheraddition of supplemental promoter is made until the start of a secondreaction run. For this purpose, a reaction run consists of the period ofoxidation of formaldehyde and formic acid from the time of any discreteaddition of supplemental promoter to the reaction zone until the time ofthe next succeeding addition of supplemental promoter to the reactionzone, and may consist of 50 or more, typically 100 or more, turnovers ofthe working volume of the reactor.

As noted, only a fraction of the supplemental promoter added to thefirst batch of a cycle remains on the catalyst after multiple cycles ofa series of batch reaction runs, or after multiple turnovers of acontinuous reaction run. However, the supplemental promoter remainseffective to enhance the oxidation of theN-(phosphonomethyl)iminodiacetic acid substrate, formaldehyde and formicacid when the substrate is contacted with the oxidizing agent in areaction zone which comprises the liquid reaction medium and wherein themass ratio of supplemental promoter to the catalyst in such reactionzone is at least about 1:200,000, preferably at least about 1:70,000,more preferably at least about 1:30,000, most preferably at least about1:15,000. Inasmuch as substantially the sole source of supplementalpromoter for the reactor may be recycle catalyst, it is furtherpreferred that the supplemental promoter be present on or in the recyclecatalyst in the same mass ratios, i.e., at least about 1:200,000,preferably at least about 1:70,000, more preferably at least about1:30,000, most preferably at least about 1:15,000.

The supplemental promoter content of the reaction zone can also beexpressed as a mass ratio to the noble metal component of the catalyst.For example, for a 5% noble metal on carbon catalyst, the ratio ofsupplemental promoter to noble metal should be at least about 1:10,000,more preferably 1:3500, more preferably 1:1800, most preferably 1:700.These preferences generally prevail over the range of noble metalcontent of the noble metal on carbon catalyst, which is typically fromabout 0.5 to 20% noble metal. However, where the noble metal content isrelatively high, e.g., approaching 20%, the supplemental promoter may beeffective in relatively lower mass ratios to the noble metal component,even as low as 1:40,000.

Where the supplemental promoter is added in a discrete charge at thestart of a series of batch reaction cycles, or at the beginning of acontinuous reaction run as defined above, it is added in a mass ratio tothe noble metal component of the catalyst of at least about 1:750,preferably at least about 1:250, more preferably at least about 1:125,most preferably at least about 1:50. As indicated above, the preferredratio of supplemental promoter to noble metal may vary with the noblemetal content of the catalyst. Thus, e.g., when the noble metal contentof the catalyst approaches 20% by weight, the supplemental promoter maybe effective when added at a mass ratio to noble metal of 1:3000 orhigher, more preferably at least about 1:1000, 1:500 or 1:200.

Periodic discrete additions of supplemental promoter may be advantageousbecause excessive proportions of supplemental promoter, while maximizingthe effectiveness of the catalyst for the oxidation of formaldehyde andformic acid, may retard the oxidation of theN-(phosphonomethyl)iminodiacetic acid substrate. By adding supplementalpromoter only periodically, the proportions of supplemental promoterdeposited on the catalyst and present in the reaction zone may decayfairly rapidly to a residual quasi-steady state range wherein thesupplemental promoter remains effective to enhance catalytic activityfor the oxidation of formaldehyde or formic acid without significantlyretarding the rate or extent of oxidation of theN-(phosphonomethyl)iminodiacetic acid substrate. Thus, the optimumsupplemental promoter content within the oxidation reaction zone foroxidizing the N-(phosphonomethyl)iminodiacetic acid substrate, and onthe recycle catalyst for such reaction, may be lower than 1:15,000, forexample, in a range of 1:65,000 to 1:25,000.

Deposit of supplemental promoter on the surface of a noble metal oncarbon catalyst in the reaction medium results in formation of a novelcatalyst complex comprising the catalyst and the promoter. The catalystcomponent of the catalyst complex may further comprise a surfacepromoter comprising a metal different from the supplemental promoter or,in some instances, comprising the same metal. The supplemental promoteris believed to be deposited by adsorption from the reaction medium, andremains desorbable from the catalyst surface into the catalyst medium.While an operative fraction of residual supplemental promoter resistsdesorption to remain adhered to the catalyst through multiple reactioncycles (or through an extended run of a continuous reaction system) asexplained hereinabove, the supplemental promoter is typically moredesorbable than the surface promoter which is applied in the catalystpreparation process.

As described above, the catalyst is prepared in the first instance bydepositing noble metal and optionally surface promoter onto a carbonsupport to form a catalyst precursor, then reducing the catalystprecursor to produce the reaction catalyst. The novel catalyst complexis formed by subsequent deposition of supplemental promoter on theoxidation catalyst, typically by adsorption to the carbon or noble metalsurface. Advantageously, the supplemental promoter is mixed with theoxidation catalyst in the reaction medium so that the promoter isdeposited from the reaction medium onto the catalyst surface. However,it will be understood that, in the alternative, the supplementalpromoter can be premixed with the oxidation catalyst in another liquidmedium to form the catalyst complex, after which the catalyst complexmay be introduced into the reaction medium for use in conducting theoxidation reaction.

It should be recognized that, depending on the desired effects, morethan one supplemental promoter may be used. In addition, eachsupplemental promoter may come from more than one source. Further, thecarbon-supported, noble-metal-containing catalyst may already contain anamount of metal on its surface which is the same metal as thesupplemental promoter, such as where (a) the catalyst is manufacturedwith a such a metal on its surface to act as a catalyst-surfacepromoter, or (b) the catalyst is a used catalyst which has beenrecovered from a previous reaction mixture where the metal was present(e.g., as a supplemental promoter).

In a particularly preferred embodiment, the carbon-supported,noble-metal-containing catalyst itself also comprises one or morecatalyst-surface promoters on its surface, as described above. Where thecatalyst is being used in the oxidation of anN-(phosphonomethyl)iminodiacetic acid substrate and the supplementalpromoter is bismuth, it is particularly preferred for the catalyst tocontain tin and/or iron (the presence of tin tends to be particularlyuseful for increasing the oxidation of the formaldehyde by-product inaddition to increasing the oxidation of the formic acid by-product).

In many instances, after a supplemental promoter (particularly bismuth)and a carbon-supported, noble-metal-containing catalyst have beencombined, at least a portion of the supplemental promoter deposits ontothe surface of the carbon support and/or noble metal of the catalyst,and is consequently retained by the catalyst. Because the catalystretains the promoter, the catalyst may typically be recycled for use incatalyzing the oxidation of subsequent amounts of the oxidationsubstrate (e.g., the catalyst may be used to oxidize additional batchesof the oxidation substrate, or may be used in a continuous oxidationprocess) while still retaining the benefits of the supplementalpromoter. And, as the effects of the supplemental promoter decrease overtime with use, replenishing amounts of fresh supplemental promoter mayperiodically be mixed with the catalyst to revive the effects and/orachieve other desired results (e.g., decreased formic acid levels).Where, for example, the catalyst is used in multiple batch reactions,such periodic replenishing may, for example, be conducted after thecatalyst has been used in at least about 20 batch oxidation reactions(more preferably after it has been used in at least about 30 batchoxidation reactions, and most preferably after it has been used in atleast about 100 or more batch oxidation reactions). Where a catalyst isperiodically replenished with fresh supplemental promoter, the mixingfor replenishment may be conducted in or separately from the oxidationreaction zone(s).

In a particularly preferred embodiment, a supplemental promoter is mixedwith a used catalyst (i.e., a catalyst that has been used in one or moreprevious oxidation reactions). Typically, the activity and/or desiredselectivity of a catalyst decreases with use. Thus, for example, theactivity of a carbon-supported, noble-metal-containing catalyst foroxidizing by-products (e.g., formaldehyde and/or formic acid) of theoxidation of an N-(phosphonomethyl)iminodiacetic acid substrate oftentends to decrease as the catalyst is used, thereby causing lessformaldehyde and/or formic acid to be destroyed and, consequently, agreater amount of N-methyl-N-(phosphonomethyl)glycine to be produced.Eventually, in fact, this activity will decrease to a level where anunacceptable amount of formaldehyde and/or formic acid is not oxidized,consequently often causing an unacceptable amount ofN-methyl-N-(phosphonomethyl)glycine compounds to be produced (i.e., theselectivity of the catalyst for making N-(phosphonomethyl)glycinecompounds from N-(phosphonomethyl)iminodiacetic acid substrates willdecrease to an unacceptable level). Traditionally, when the catalystactivity for oxidizing the by-products reaches such a point, thecatalyst has been deemed unuseable, and, consequently, has either beenrecycled (i.e., reactivated) through a time-consuming and sometimescostly process, or discarded altogether. It has been discovered inaccordance with this invention, however, that such a catalyst can berevived (i.e., the selectivity of the catalyst for makingN-(phosphonomethyl)glycine product can be increased to an acceptablelevel) by mixing the catalyst with a supplemental promoter, particularlybismuth or tellurium. In other words, the supplemental promoter can beused to modify the catalyst performance and extend the life of thecatalyst.

It has been observed that a supplemental promoter (particularly bismuth)may cause a slight decrease in the oxidation rate of anN-(phosphonomethyl)iminodiacetic acid substrate. In such an instance,the oxidation rate may typically be increased, at least in part, byincreasing the amount of oxygen fed into the reacting mixture,maintaining a relatively high oxygen flowrate for an extended periodduring the reaction, and/or increasing the pressure. Where, however, theoxygen flow is increased, it preferably is not increased to an extentwhich causes the catalyst surface to become detrimentally over-oxidized.Thus, the increased oxygen feed rate preferably is maintained at a levelsuch that at least about 40% (more preferably at least about 60%, evenmore preferably at least about 80%, and most preferably at least about90%) of the fed oxygen is utilized.

Preferred Oxidation Reactor Systems

The oxidation reaction zone(s) may comprise a wide range of reactorconfigurations, including those that have back-mixed characteristics, inthe liquid phase and optionally in the gas phase as well, and those thathave plug flow characteristics. Suitable reactor configurations havingback-mixed characteristics include, for example, stirred tank reactors,ejector nozzle loop reactors (also known as venturi-loop reactors) andfluidized bed reactors. Suitable reactor configurations having plug flowcharacteristics include those having a packed or fixed catalyst bed(e.g., trickle bed reactors and packed bubble column reactors) andbubble slurry column reactors. Fluidized bed reactors may also beoperated in a manner exhibiting plug flow characteristics.

In a broad sense, the oxidation reaction may be practiced in accordancewith the present invention at a wide range of temperatures, and atpressures ranging from sub-atmospheric to super-atmospheric. Use of mildconditions (e.g., room temperature and atmospheric pressure) hascommercial advantages in that less expensive equipment may be used inthe reactor system. However, operating at higher temperatures andsuper-atmospheric pressures, tends to improve mass transfer between theliquid and gas phase (e.g., the oxygen source), increase theN-(phosphonomethyl)iminodiacetic acid oxidation reaction rate andincrease the N-(phosphonomethyl)glycine product solubility, therebyreducing the amount of water requiring separation to precipitate andrecover the product. Accordingly, use of more aggressive oxidationconditions can actually reduce the overall costs of a plant and reduceoperating costs per unit of N-(phosphonomethyl)glycine product produced.Preferably, the N-(phosphonomethyl)iminodiacetic acid oxidation reactionis conducted at a temperature of from about 20° C. to about 180° C.,more preferably from about 50° C. to about 140° C., still morepreferably from about 80° C. to about 110° C., and yet still morepreferably from about 95° C. to about 105° C. At temperatures greaterthan about 180° C., the feed materials tend to slowly decompose.Moreover, the selectivity toward the desired N-(phosphonomethyl)glycineproduct tends to worsen as the oxidation reaction temperature increasesmuch above about 90° C. For example, the production of the undesiredby-product-methyl-N-(phosphonomethyl)glycine (NMG) tends to increaseroughly 2- to 4-fold for each 10° C. increase in reaction temperatureabove 90° C. Lower temperatures (i.e., temperatures less than about 95°C.) often tend to be less advantageous because the solubility of someN-(phosphonomethyl)iminodiacetic acid substrates andN-(phosphonomethyl)glycine products are reduced at such temperatures.The total pressure maintained in the oxidation reaction zone(s)generally depends on the temperature used and the reactor configuration.The total pressure in each oxidation reaction zone is preferably atleast equal to atmospheric pressure and sufficient to prevent the liquidreaction medium in the oxidation zone from boiling. Preferred oxidationreaction conditions for particular reactor systems are discussed ingreater detail below.

In a preferred embodiment, the oxidation of theN-(phosphonomethyl)iminodiacetic acid substrate is conducted in one ormore continuous oxidation reaction zones, wherein the substrate iscontinuously oxidized to form the N-(phosphonomethyl)glycine product.Continuous oxidation provides the opportunity for greater processthroughput and lower production costs. Moreover, because the oxidationof the N-(phosphonomethyl)iminodiacetic acid substrate is exothermic,after startup of a continuous oxidation reactor system, typically noheat input to the aqueous feed stream introduced into the oxidation zoneis required to maintain the desired oxidation reaction temperature.

Various reactor configurations may be suitably employed to provide thecontinuous oxidation reaction zone(s). In accordance with one preferredembodiment, continuous oxidation of the N-(phosphonomethyl)iminodiaceticacid substrate is carried out in one or more substantially back-mixedoxidation reaction zones (i.e., back-mixed in at least the liquid phase)utilizing a heterogeneous particulate catalyst, preferably the deeplyreduced noble metal on carbon particulate catalyst described above,suspended in contact with the liquid reaction medium. However, it shouldbe understood that the practice of the present invention is not limitedto use of such a deeply reduced catalyst, nor to a catalyst inparticulate form. Moreover, it should be understood that the catalystused in the reactor systems of the present invention may comprise amixture of different catalysts and/or the catalyst may vary from oneoxidation reaction zone to the next within the reactor system.

FIG. 2 shows an example of a reactor system that may be used to carryout the continuous oxidation process of the present invention. Thesystem shown in FIG. 2 comprises a continuous stirred tank reactor 3providing mechanical agitation of the liquid reaction medium containedtherein, typically by a rotating impeller. Stirred tank reactorssuitably back-mix the liquid phase within the reaction zone, arerelatively simple in design, and operation can be scaled to the desiredprocess capacity. Various impeller designs may be employed, includingsystems with multiple blades rotated on a common shaft. The reactorvessel may include internal baffles and/or draft tubes to modify mixingcharacteristics and prevent swirling of the liquid reaction medium as iswell-known to those skilled in the art.

Although the reactor system shown in FIG. 2 comprises a singlecontinuous stirred tank reactor, in many instances, a reactor systemcomprising two or more back-mixed oxidation reaction zones staged inseries is preferred as will be described in greater detail below. Theback-mixed oxidation reaction zone(s) may be suitably provided byreactor configurations other than continuous stirred tank reactors(e.g., ejector nozzle loop reactors and fluidized bed reactors).Moreover, different reactor configurations may be combined in a reactorsystem comprising multiple oxidation reaction zones. For example, one ormore reactors having back-mixed characteristics may be combined with areactor configuration having plug flow characteristics such as a fixedcatalyst bed reactor.

An aqueous feed stream 1 comprising the N-(phosphonomethyl)iminodiaceticacid substrate is continuously or intermittently introduced into aliquid reaction medium within the stirred tank reactor 3 along with anoxygen source. The heterogeneous particulate catalyst is also presentwithin the oxidation reaction zone in contact with the liquid reactionmedium and used to catalyze the oxidation of theN-(phosphonomethyl)iminodiacetic acid substrate in the aqueous feed.Vapor comprising CO₂ evolved as the oxidation reaction proceeds isvented from the headspace above the reaction mixture in the stirred tankreactor 3. A reaction mixture effluent 7 containing theN-(phosphonomethyl)glycine product and the heterogenous particulatecatalyst is continuously or intermittently withdrawn from the stirredtank reactor 3 and transferred to a catalyst filter 9, whereinsubstantially all the catalyst is separated from the reaction mixture toform: (1) a catalyst recycle stream 11 comprising substantially all thecatalyst and a residual amount of the N-(phosphonomethyl)glycineproduct; and (2) a filtrate 13 containing much of theN-(phosphonomethyl)glycine product. The catalyst recycle stream 11 isreintroduced into the stirred tank reactor 3, while the filtrate 13 iscarried forward to concentrate and purify the N-(phosphonomethyl)glycineproduct.

The temperature within the oxidation reaction zone is preferablymaintained sufficiently high with respect to theN-(phosphonomethyl)glycine product concentration such that essentiallyall the N-(phosphonomethyl)glycine product in the reaction mixtureeffluent 7 withdrawn from the stirred tank reactor 3 is dissolved. Thus,for example, when the N-(phosphonomethyl)glycine product is theN-(phosphonomethyl)glycine free acid at a concentration of from about 7to about 15% by weight, the temperature of the reaction mixture effluentwithdrawn from the stirred tank reactor 3 is preferably maintained atfrom about 80° C. to about 180° C., more preferably from about 90° C. toabout 150° C., more preferably from about 90° C. to about 135° C., evenmore preferably from about 95° C. to about 125° C., and still morepreferably from about 95° C. to about 105° C. However, it should beunderstood that precipitation of the N-(phosphonomethyl)glycine productin the reaction mixture effluent 7 can be tolerated and satisfactoryresults still obtained. The precipitated N-(phosphonomethyl)glycineproduct may be separated with the particulate catalyst (e.g.,co-filtered) from the remainder of the reaction mixture effluent 7.

At startup, the reaction mixture in the oxidation reaction zone 3 and/orthe aqueous feed stream 1 may be heated to obtain the desiredtemperature for the oxidation reaction. If heat addition is required,all or at least a portion of the heat energy may be provided in pumpingthe various feed streams into the stirred tank reactor 3 and through theremainder of the reactor system such that a separate conventional feedpreheater may not be necessary. Because the oxidation reaction isexothermic, it will normally be necessary to remove heat energy from thereaction mixture once the oxidation reaction begins to evolvesignificant amounts of heat in order to maintain the desired temperaturewithin the oxidation zone. As shown in FIG. 2, excess reaction heat maybe extracted from the reaction mixture within the stirred tank reactor 3by passing the reaction mixture through an external heat transferrecirculation loop 15 containing a heat exchanger 16 wherein heat istransferred indirectly from the reaction mixture to a cooling medium(e.g., cooling water). The reaction temperature is controlled by, forexample, controlling the supply of cooling water to heat exchanger 16 inresponse to the signal from a temperature controller. Reaction heat canbe removed from the oxidation reaction zone by other conventional meansas well, such as with cooling coils immersed within the reaction mixtureor a reactor vessel jacket through which a cooling medium is circulated.

The total pressure in the stirred tank reactor 3 is generally from about0 to about 500 psig and is preferably maintained sufficiently high toprevent the liquid reaction medium therein from boiling. Typically, thetotal pressure in the stirred tank reactor 3 is from about 30 to about500 psig. When operating the oxidation reaction zone within theespecially preferred temperature range of from about 95° C. to about105° C., the total pressure maintained within the stirred tank reactor 3is preferably from about 30 to about 130 psig and more preferably fromabout 90 to about 110 psig.

A wide range of N-(phosphonomethyl)iminodiacetic acid substrateconcentrations may be used in the aqueous feed stream 1. The aqueousfeed stream 1 includes the catalyst recycle stream 11 and any otherrecycle streams from other parts of the process introduced into stirredtank reactor 3. In slurry reactors, such as the stirred tank reactorshown in FIG. 2, the substrate concentration in the aqueous feed stream1 is preferably selected with respect to the temperature of the reactionmixture effluent 7 such that essentially all of the desiredN-(phosphonomethyl)glycine product is dissolved. As noted above,substrate concentrations which form reaction mixtures containingN-(phosphonomethyl)glycine product at a concentration exceeding thesolubility limit of the product may also be employed, but are generallyless preferred. Relative to many commonly-practiced commercialprocesses, this invention allows for greater temperatures andN-(phosphonomethyl)iminodiacetic acid substrate concentrations to beused to prepare N-(phosphonomethyl)glycine while minimizing by-productformation. In the commonly practiced commercial processes using acarbon-only catalyst, it has often been economically preferable tooperate at low substrate concentrations and temperatures to minimize theformation of the N-methyl-N-(phosphonomethyl)glycine by-product. Withthose processes and catalysts, temperatures of from about 60° C. to 90°C. are typically used to achieve cost effective yields and to minimizethe generation of waste. At such temperatures, the maximumN-(phosphonomethyl)glycine solubility typically is less than 6.5% ([massof N-(phosphonomethyl)iminodiacetic acid substrate÷total reactionmass]×100%). But, with the preferred oxidation catalyst and reactionprocess of this invention, the loss of noble metal from the catalyst andcatalyst deactivation is minimized and the formaldehyde is moreeffectively oxidized, thereby allowing for reaction temperatures asgreat as 180° C. (or greater). The use of higher oxidation reactiontemperatures and reactor feed concentrations permits reactor throughputto be increased, reduces the amount of water that must be removed perpound of N-(phosphonomethyl)glycine product produced, and reduces thecost of manufacturing N-(phosphonomethyl)glycine. This invention thusprovides economic benefits over many commonly-practiced commercialprocesses.

The preferred upper limit on the concentration of theN-(phosphonomethyl)iminodiacetic acid substrate is dependent on thespecific substrate employed. For example, in the case of a salt ofN-(phosphonomethyl)iminodiacetic acid (e.g., potassium salt)concentrations up to about 70 wt. % may be employed. Typically, however,an N-(phosphonomethyl)iminodiacetic acid substrate concentration of upto about 50 wt. % is preferred (especially at a reaction temperature offrom about 20 to about 180° C.). More preferably, anN-(phosphonomethyl)iminodiacetic acid substrate concentration of up toabout 25 wt. % is used (particularly at a reaction temperature of fromabout 60 to about 150° C.). Even more preferably, anN-(phosphonomethyl)iminodiacetic acid substrate concentration of fromabout 3 to about 20 wt. % is used (particularly at a reactiontemperature of from about 100 to about 130° C.). At preferred reactiontemperatures of from about 95 to about 105° C., theN-(phosphonomethyl)iminodiacetic acid substrate concentration preferablyis from about 7 to about 15 wt. %, more preferably from about 7 to about12% by weight, and even more preferably from about 9 to about 10 wt. %.

In some instances, the source of the N-(phosphonomethyl)iminodiaceticacid substrate fed to the process in the aqueous feed stream 1 containschloride ions (Cl⁻) which have been carried over from the synthesis ofthe substrate. Where the catalyst comprises a carbon-supported noblemetal, chloride ions tend to interact with the catalyst to increaseleaching of the noble metal and inhibit formic acid by-productoxidation. Moreover, chloride levels may tend to elevate in reactorsystems in which streams (e.g., from the product concentrating and/orpurifying steps of the process) are recycled and introduced into theoxidation reaction zone(s) as described below. Preferably, the chlorideion concentration in the liquid phase reaction medium in contact withthe catalyst within the oxidation reaction zone(s) is maintained at nogreater than about 500 ppm by weight, more preferably no greater thanabout 300 ppm by weight, and even more preferably no greater than 100ppm by weight. Advantageously, control of chloride levels in theoxidation reaction zone(s) is established by using anN-(phosphonomethyl)iminodiacetic acid substrate source having arelatively low chloride content to form the aqueous feed stream 1.Preferably, the concentration of chloride ion in the source of theN-(phosphonomethyl)iminodiacetic acid substrate fed to the process inthe aqueous feed stream is less than about 5000 ppm by weight, morepreferably less than about 3000 ppm by weight, even more preferably lessthan about 2000 ppm by weight, and especially less than about 1000 ppmby weight on a dry basis. An N-(phosphonomethyl)iminodiacetic acidsubstrate source meeting such standards may be produced, for example, bythe processes described in U.S. Pat. Nos. 4,724,103 and 4775,498, whichare expressly incorporated herein by reference. In addition, it may beadvantageous to utilize a deeply reduced noble metal (e.g., platinum) oncarbon catalyst modified with an addition of ruthenium as describedabove to catalyze the reactions in a continuous oxidation reactorsystem. Such ruthenium modified catalysts may provide increasedresistance to chloride attack and noble metal leaching and may beparticularly suited for use in a continuous oxidation reactor systemwhere chloride levels in the oxidation reaction zone(s) are elevated dueto various recycle streams.

The oxygen source may be introduced into the reaction mixture within thestirred tank reactor 3 by any conventional manner which maintains thedissolved oxygen concentration in the reaction mixture at the desiredlevel. Preferably, the oxygen source is an O₂-containing gas such asair, pure O₂ or O₂ diluted with one or more non-oxidizing gases (e.g.,He, Ar and N₂). More preferably, the oxygen source is an O₂-containinggas comprising at least about 95 mole % O₂, typically approximately 98mole % O₂. The O₂-containing gas is introduced into the reaction mixturein a manner which provides intimate contact of the gas with the reactionmixture. For example, an O₂-containing gas may be introduced through asparger conduit or similar distributor positioned in the bottom of thestirred tank reactor 3 below the impeller so that the turbulence inducedby the rotating impeller intimately mixes and distributes theO₂-containing gas as it rises though the liquid reaction medium.Distribution of the O₂-containing gas within the reaction mixture may befurther enhanced by passing the gas through a diffuser such as a porousfrit or by other means well-known to those skilled in the art.Alternatively, the O₂-containing gas may be introduced into theheadspace above the reaction mixture in the stirred tank reactor 3.

If the dissolved oxygen concentration in the reaction mixture is toogreat, the catalyst surface tends to become detrimentally oxidized,which, in turn, tends to lead to more leaching and decreasedformaldehyde activity (which, in turn, leads to moreN-methyl-N-(phosphonomethyl)glycine being produced). To avoid thisproblem, it is generally preferred to use an oxygen feed rate such thatat least about 40%, more preferably at least about 60%, even morepreferably at least about 80%, and still even more preferably at leastabout 90% of the oxygen is utilized. As used herein, the percentage ofoxygen utilized equals: (the total oxygen consumption rate÷oxygen feedrate)×100%. The term total oxygen consumption rate means the sum of: (i)the oxygen consumption rate (R_(i)) of the oxidation reaction of theN-(phosphonomethyl)iminodiacetic acid substrate to form theN-(phosphonomethyl)glycine product and formaldehyde, (ii) the oxygenconsumption rate (R_(ii)) of the oxidation reaction of formaldehyde toform formic acid, and (iii) the oxygen consumption rate (R_(iii)) of theoxidation reaction of formic acid to form carbon dioxide and water.

The oxygen partial pressure may vary in different regions of theoxidation reaction zone(s). Preferably, the oxygen partial pressure inthe headspace above the liquid reaction mixture in a stirred tankreactor is from about 0.1 to about 35 psia, more preferably from about 1to about 10 psia.

When the oxidation reaction is conducted in a single continuous stirredtank reactor system, the residence time in the reactor 3 can vary widelydepending on the specific catalyst and oxidation reaction conditionsemployed. Typically, the residence time is from about 3 to about 120minutes, more preferably from about 5 to about 90 minutes, still morepreferably from about 5 to about 60 minutes, and still even morepreferably from about 15 to about 60 minutes. The residence time isdefined relative to the flowrate of filtrate 13 and the working volumeof stirred tank reactor 3.

The particulate catalyst utilized in the continuous oxidation reactionsystem may comprise a support in the form of a powder exhibiting aparticle size distribution as previously described. Preferably, theaverage particle size of the particulate catalyst is from about 15 toabout 40 μm, more preferably about 25 μm. The concentration of theparticulate catalyst in the reaction mixture within the stirred tankreactor 3 is preferably from about 0.1 to about 10 wt. % ([mass ofcatalyst÷total reaction mass]×100%). More preferably, the catalystconcentration is from about 0.5 to about 5 wt. %, even more preferablyfrom about 1 to about 3 wt. %, and still even more preferably about 2wt. %. Concentrations greater than about 10 wt. % are difficult toseparate from the N-(phosphonomethyl)glycine product. On the other hand,concentrations less than about 0.1 wt. % tend to produce unacceptablylow reaction rates.

The catalyst filter 9 used to separate the particulate catalyst from thereaction mixture 7 withdrawn from the stirred tank reactor 3 ispreferably a filter adapted for continuous separation of catalyst fromthe reaction mixture. That is, the catalyst filter 9 is capable ofreceiving a continuous flow of reaction mixture 7 and continuouslyforming the filtrate 13 and the catalyst recycle stream 11 withouthaving to interrupt the flow of reaction mixture introduced into thefilter. In accordance with an especially preferred embodiment, catalystfilter 9 is a continuous cross-flow filter or a continuous back-pulsefilter. In practicing the continuous oxidation process depicted in FIG.2, a back-pulse filter is generally preferred over a cross-flow filterbecause the present commercially available back-pulse filters typicallycan form a catalyst recycle stream 11 containing a greater concentrationof catalyst, often at least a 5-fold greater catalyst concentration, ascompared to present commercially available cross-flow filters.

FIG. 2A is a schematic flow sheet of a continuous reactor system similarto that shown in FIG. 2 particularly adapted for use of a continuousback-pulse filter as catalyst filter 9. When the operating totalpressure in the oxidation reaction zone(s) is much higher thanatmospheric pressure, as is preferred, the pressure over the reactionmixture effluent 7 withdrawn from the stirred tank reactor 3 istypically reduced in connection with concentrating and purifying theN-(phosphonomethyl)glycine product. At least a portion of this pressurereduction may take place in a flash tank 17 upstream of catalyst filter9. The flash tank 17 lowers the pressure on the reaction mixture 7 tosome degree, causing dissolved CO₂ to be flashed out of the mixture andvented as vapor from the flash tank. Flash tank 17 reduces the pressureat which the continuous back-pulse catalyst filter 9 must operate,thereby reducing the capital costs and complexity of the filter system.An oxygen source (e.g., an O₂-containing gas) may be introduced (e.g.,sparged) into the flash tank 17 to further oxidizeN-(phosphonomethyl)iminodiacetic acid substrate in the reaction mixture7 that did not oxidize in the stirred tank reactor 3, as well as tofurther oxidize formaldehyde and formic acid by-products present in thereaction mixture. In this manner, the flash tank may 17 act as anadditional oxidation reaction zone in series with the stirred tankreactor 3.

The continuous back-pulse filter system comprises a filter element andis preferably operated adiabatically, but may be provided with heatingor cooling capability. Preferably, the liquid used to back-pulse thefilter element and remove separated catalyst is a portion of thefiltrate 13. The filtrate 13 is carried forward to concentrate andpurify the N-(phosphonomethyl)glycine product, while the recyclecatalyst stream 11 is continuously withdrawn from the catalyst filter 9and transferred to an optional catalyst holding tank 5 (also called acatalyst recycle tank or catalyst slurry tank) before the catalyst isreintroduced into stirred tank reactor 3.

Although the catalyst filter 9 in the oxidation reactor system shown inFIGS. 2 and 2A is preferably a continuous back-pulse filter, it shouldbe recognized that continuous cross-flow filters are in some instancesmore preferred. The system depicted in FIG. 2B is similar to that shownin FIGS. 2 and 2A except that the catalyst filter 9 is placed within theexternal heat transfer recirculation loop 15 rather than in a separatecatalyst recycle loop. In such an embodiment, catalyst filter 9 ispreferably a continuous cross-flow filter. Typically, a pre-filter flashtank is not employed in conjunction with a cross-flow filter.Furthermore, due to the relatively large volume of the catalyst recyclestream 11 issuing from a continuous cross-flow filter, a catalystholding tank is likewise typically omitted.

Aside from cross-flow and back-pulse filters, the catalyst filter 9 usedin a continuous oxidation reactor system may alternatively be a vacuumfilter or may comprise a bank of leaf filters used to treat a continuousflow of reaction mixture effluent 7 in staggered filtration cycles. As afurther alternative, stirred tank reactor 3 may include an internalcatalyst filter (e.g., a porous frit) which blocks the particulatecatalyst from being withdrawn with the reaction mixture effluent 7 suchthat the catalyst is substantially retained within the oxidationreaction zone and the reaction mixture effluent is substantially free ofthe particulate catalyst. Moreover, it should be recognized that othermeans of catalyst separation may be used instead of (or in addition to)the catalyst filter 9. For example, the catalyst could be separated fromthe oxidation reaction mixture effluent using a centrifuge.

As the catalyst deactivates with use, it may be at least partiallyreactivated either continuously or intermittently. Reactivation maycomprise reducing the surface of the catalyst after it has becomeheavily oxidized. In such an instance, the surface may, for example, bewashed to remove the organics, and then reduced using the reductiontreatments described above. Such a reducing treatment may comprise, forexample, continuous or intermittent introduction of a reducing agentinto the reactor system. For example, the reducing agent may compriseformaldehyde and/or formic acid and may often advantageously be obtainedfrom various recycle streams described herein. Reactivation may also beachieved by, for example, introducing a supplemental promoter,especially bismuth oxide into the reactor system as described above. Inaccordance with a preferred embodiment of the present invention, asupplemental promoter (e.g., Bi₂O₃) is introduced continuously orintermittently into the continuous reactor system such that theconcentration of formic acid in the reaction mixture effluent withdrawnfrom the last oxidation reaction zone is maintained at less than about6000 ppm, more preferably from about 1000 ppm to about 3000 ppm. Inaccordance with an especially preferred practice of the presentinvention, the concentration of formic acid in the reaction mixtureeffluent withdrawn from the last oxidation reaction zone is monitored.Once the measured concentration exceeds about 6000 ppm, more preferablyabout 3000 ppm, even more preferably about 2000 ppm, continuous orintermittent introduction of a supplemental promoter into to the reactorsystem is initiated and continued until the concentration of formic acidin the reaction mixture effluent withdrawn from the last oxidationreaction zone begins to decline. Preferably, the rate of addition of thesupplemental promoter to the reactor system is such that theconcentration of formic acid in the reaction mixture effluent withdrawnfrom the last oxidation reaction zone continues to rise for a period oftime after addition of a supplemental promoter to the system hascommenced. In the case of Bi₂O₃ added to the reactor system as asupplemental promoter, the weight ratio of Bi₂O₃ to theN-(phosphonomethyl)iminodiacetic acid substrate fed to the system isfrom about 1:20,000,000 to about 1:200,000.

Although optional in the continuous oxidation reactor system shown inFIG. 2A, the catalyst holding tank 5 may be advantageous when a deeplyreduced particulate catalyst is used because it provides a place for thecatalyst mass to be uniformly reactivated. As shown in FIG. 2A, areducing agent 18 and/or a supplemental promoter 19 may be introducedinto the catalyst holding tank 5 containing recycled catalyst. Thereducing agent and/or supplemental promoter may alternatively be addeddirectly to the oxidation reaction zone(s) or introduced elsewhere intothe reactor system. It should be further recognized that merely allowingthe recycled catalyst to sit in the catalyst holding tank 5 with theresidual reaction mixture may also beneficially reduce the catalystsurface (particularly a catalyst comprising carbon-supported noblemetal). Preferably, the catalyst holding tank is substantially free ofO₂ and other oxidizing gases. Accordingly, it may be advantageous tointroduce (e.g., sparge) nitrogen or other non-oxidizing gas into thetank 5 to help remove O₂. Allowing the slurry of particulate catalystand residual slurry to remain outside the oxidation reaction zone(s) inan environment substantially free of O₂ for a period of time beforebeing reintroduced into the oxidation reaction zone(s) is believed toreduce the surface of the catalyst and achieve a degree of reactivationand extend the useful life of the catalyst. The catalyst holding tank orcatalyst slurry tank 5 may have various configurations, but is typicallya stirred tank in which the catalyst slurry comprising the particulatecatalyst and residual reaction mixture is agitated with a rotatingimpeller to improve uniformity in the catalyst slurry by preventing thecatalyst from settling to the bottom of the tank 5 and promote uniformreactivation of the catalyst as well. The residence time of the catalystin the catalyst holding tank 5 may be adjusted by adjusting the catalystslurry volume in the catalyst holding tank relative to the workingvolume of reaction medium within the oxidation reaction zone(s). Longercatalyst residence times in the catalyst holding tank 5 are generallybeneficial to catalyst performance. However, since longer residencetimes require a larger catalyst inventory in the reactor system, thebenefits of longer residence times must be weighed against the increasedcatalyst costs, which may become significant, especially in the case ofa catalyst comprising a carbon-supported noble metal. Preferably, theresidence time of the recycled catalyst in the catalyst holding tank isat least about 2 minutes, more preferably at least about 5 minutes, evenmore preferably from about 5 to about 40 minutes.

Reduced losses of noble metal may be observed with this invention if asacrificial reducing agent is maintained or introduced into the reactionsolution. Suitable reducing agents include formaldehyde, formic acid,and acetaldehyde. Most preferably, formic acid, formaldehyde, ormixtures thereof (e.g., obtained from the various recycle streamsdescribed herein) are used.

Catalyst (e.g., catalyst having diminished activity and/or selectivity)may also be continuously or intermittently purged from the continuousoxidation reactor system via catalyst purge stream 20, and replaced withfresh catalyst via the fresh catalyst feed stream 21. Whenintermittently purging the catalyst, the entire catalyst mass may bepurged from the process at the same time (which is typically the morepreferred method), or a fraction of the catalyst mass may be purged atvarious time increments. In other words, intermittent purging includesany repeated purging of catalyst that is not continuous.

In accordance with a more preferred embodiment of the present invention,the continuous oxidation of an N-(phosphonomethyl)iminodiacetic acidsubstrate in the presence of a particulate heterogenous catalyst slurryis staged in two or more substantially back-mixed oxidation reactionzones (i.e., back-mixed in at least the liquid phase) operated inseries. A combination of two or more back-mixed oxidation reaction zonesin series is advantageous because such a reactor system tends to behavemore like a plug flow reactor, producing fewer by-products and improvingthe yield of the N-(phosphonomethyl)glycine product. Moreover, thecombination of two or more reaction zones provides the ability to varyreaction conditions in accord with the prevailing reaction kinetics atdifferent stages of the oxidation reaction. The second and subsequentoxidation reaction zone(s) may provide further conversion ofN-(phosphonomethyl)iminodiacetic acid substrate and/or oxidation of C₁by-products (e.g., formaldehyde and formic acid).

The oxidation of an N-(phosphonomethyl)iminodiacetic acid substratebehaves approximately as a zero order reaction with respect to thesubstrate concentration until the N-(phosphonomethyl)iminodiacetic acidsubstrate concentration decreases to no greater than about 4.5% byweight, more typically to no greater than about 2.7% by weight, evenmore typically from about 0.4% to about 1.8% by weight, still even moretypically from about 0.4% to about 1.3% by weight, and still yet evenmore typically no greater than about 1% by weight. Where, for example,the substrate concentration in the aqueous feed to the first reactionzone is about 9% by weight, the reaction will tend to behaveapproximately as a zero order reaction with respect to the substrateuntil at least about 50%, more typically at least about 70%, even moretypically from about 80% to about 95%, and still even more typicallyfrom about 85% to about 95% of the substrate has been consumed. At thatpoint, the oxidation rate becomes a stronger function of the substrateconcentration (i.e., the oxidation approaches first order behavior withrespect to the substrate concentration), and consequently tends todecrease as the substrate concentration further decreases. As theoxidation rate becomes a stronger function of theN-(phosphonomethyl)iminodiacetic acid substrate concentration, theoxidation of the substrate tends to be slower than the simultaneousoxidation reactions of the formaldehyde and formic acid by-products.

By utilizing a continuous oxidation reactor system comprising two ormore oxidation reaction zones in series, the residence time and/oroxygen feed in the first reaction zone may be controlled so that thereaction in the first reaction zone substantially behaves as azero-order reaction with respect to the substrate concentration (i.e.,the residence time in the first reactor may be controlled so that theconversion of substrate in the first reactor is sufficient to form areaction mixture having a substrate concentration of no greater thanabout 4.5% by weight, more preferably no greater than about 2.7% byweight, even more preferably from about 0.4% to about 1.8% by weight,still even more preferably from about 0.4% to about 1.3% by weight, andstill yet even more preferably about 1% by weight). This reactionmixture may then be transferred to the second and any subsequentreaction zones, wherein the reaction behaves substantially as afirst-order reaction with respect to the substrate concentration. Inthis manner, the reactor configuration and/or reaction conditions (e.g.,catalyst type, average catalyst age, catalyst concentration, oxygenconcentration, temperature, pressure, etc.) can be precisely controlledindependently in each reaction zone to optimize the stages of thereaction and the oxidation of the formaldehyde and formic acidby-products.

FIG. 3 shows a preferred continuous oxidation reactor system inaccordance with the present invention comprising two back-mixedoxidation reaction zones staged in series. The back-mixed oxidationreaction zones are preferably provided by two continuous stirred tankreactors 3 and 40. An aqueous feed stream 1 containing anN-(phosphonomethyl)iminodiacetic acid substrate is continuously orintermittently introduced into the first stirred tank reactor 3 alongwith an oxygen source, preferably an O₂-containing gas. TheN-(phosphonomethyl)iminodiacetic acid substrate is continuously oxidizedin the first stirred tank reactor 3 in the presence of the heterogeneousparticulate catalyst to form an intermediate aqueous reaction mixture 41comprising an N-(phosphonomethyl)glycine product and unreactedN-(phosphonomethyl)iminodiacetic acid substrate which is continuously orintermittently withdrawn from the first stirred tank reactor 3. Anintermediate aqueous feed stream 42 comprising (a)N-(phosphonomethyl)glycine product from the intermediate aqueousreaction mixture 41; and (b) unreacted N-(phosphonomethyl)iminodiaceticacid substrate, which is also, at least in part, from the intermediateaqueous reaction mixture 41, is then introduced into the second stirredtank reactor 40. Typically, additional oxygen is also introduced intothe second stirred tank reactor 40, preferably also in the form of anO₂-containing gas. In the second stirred tank reactor 40, additionalN-(phosphonomethyl)iminodiacetic acid substrate is continuously oxidizedin the presence of the heterogeneous particulate catalyst to form afinal reaction mixture effluent 45 comprising N-(phosphonomethyl)glycineproduct. The headspace above the reaction mixture within the stirredtank reactors 3 and 40 is vented to remove vapor comprising CO₂ from theoxidation reaction zones as the oxidation reaction proceeds.

Although the intermediate aqueous feed stream 42 is shown in FIG. 3 ascomprising the entire intermediate aqueous reaction mixture 41, itshould be recognized that in some embodiments of the present invention,the intermediate aqueous feed stream 42 will contain less than theentire intermediate aqueous reaction mixture 41. For example, theparticulate heterogenous catalyst may be partially or entirely removedfrom the intermediate aqueous reaction mixture 41, as described below(FIGS. 5 and 6). Furthermore, it should be understood that the first andsecond oxidation reaction zones do not have to be contained withinseparate stirred tank reactor vessels 3 and 40 as shown in FIG. 3.Multiple oxidation reaction zones may be staged in series and containedwithin a single reactor vessel divided into compartments or providedwith baffles or other means for separating one reaction zone fromanother.

In the embodiment shown in FIG. 3, the particulate catalyst flows fromthe reaction zone in the first stirred tank reactor 3 to the reactionzone in the second stirred tank reactor 40. Preferably, the particulatecatalyst is the above-described deeply reduced oxidation catalyst. Thecatalyst is continuously or intermittently introduced into the firststirred tank reactor 3 via catalyst feed stream 39. As shown in FIG. 3,the catalyst feed stream 39 is part of the aqueous feed stream 1containing an N-(phosphonomethyl)iminodiacetic acid substrate. Catalystis continuously or intermittently withdrawn from the first stirred tankreactor 3 as part of the intermediate aqueous reaction mixture 41,continuously or intermittently introduced into the second stirred tankreactor 40 as part of the intermediate aqueous feed stream 42 andfinally continuously or intermittently withdrawn from the second stirredtank reactor 40 as part of the final reaction mixture effluent 45. Thefinal reaction mixture effluent 45 is optionally depressurized in flashtank 17 and transferred to catalyst filter 9. In the catalyst filter 9,substantially all of the particulate catalyst is separated from thefinal reaction mixture effluent 45 to form (1) a catalyst recycle stream11 comprising essentially all the catalyst and a residual amount ofN-(phosphonomethyl)glycine product from the final reaction mixture 45;and (2) a filtrate 13 comprising the bulk of N-(phosphonomethyl)glycineproduct from the final reaction mixture 45. In the embodiment shown inFIG. 3, the catalyst filter 9 is preferably a continuous back-pulsefilter system in order to minimize the volume of the catalyst recyclestream and preserve the staging effect in the reactor system. Thecatalyst recycle stream 11 is directed to the catalyst holding tank 5and reintroduced into the first stirred tank reactor 3 via catalyst feedstream 39, while the filtrate 13 is carried forward to concentrate andpurify the N-(phosphonomethyl)glycine product. As the catalystdeactivates with use, it may be at least partially reactivated asdescribed above by continuously or intermittently contacting theparticulate catalyst with a reducing agent 18 (e.g., in the catalystholding tank 5) and/or introducing a supplemental promoter 19 into theprocess (e.g., into the catalyst holding tank 5 and/or directly into thefirst and/or second stirred tank reactors 3 and 40). Catalyst may becontinuously or intermittently purged from the system through thecatalyst purge stream 20 and replenished with fresh catalyst throughcatalyst feed stream 21.

During startup the reactor system in FIG. 3, the catalyst feed stream 39and/or the aqueous feed stream 1 introduced to the first stirred tankreactor 3 may be heated to obtain the desired temperature in theoxidation reaction zones. During steady state or quasi-steady stateoperations, exothermic reaction heat is ordinarily sufficient to bringfeed materials to the desired reaction temperature, and excess reactionheat is removed from the liquid reaction medium in the first reactor 3via a heat exchanger 16 in external heat transfer recirculation loop 15.The reaction temperature is controlled by, for example, controlling thesupply of cooling water to heat exchanger 15 in response to the signalfrom a temperature controller. Similarly, the temperature of the liquidreaction medium in the second oxidation reaction zone in reactor 40 maybe controlled by the rate of heat removal via heat exchanger 48 in theexternal heat transfer recirculation loop 47 associated with the secondreactor. However, the second oxidation reaction zone may be operatedwithout the heat transfer loop 47 or other means for removing reactionheat (i.e., operated adiabatically). For example, in some instances, theincremental conversion of the N-(phosphonomethyl)iminodiacetic acidsubstrate and the extant oxidation of formaldehyde and formic acid areso limited in the second stirred tank reactor 40 that the heat evolvedfrom the oxidation reactions does not necessitate cooling of thereaction mixture. Where it is desired to complete the reaction in thesecond reactor 40 at a temperature higher than the temperatureprevailing in the first reactor 3, the autogenous heat of reaction inthe second reactor may contribute all or part of the heat necessary toraise the temperature of the aqueous feed stream 42 and maintain thedesired difference in temperature between the first reactor and thesecond reactor.

The temperature of the reaction medium within the second stirred tankreactor 40 is preferably maintained high enough with respect to theN-(phosphonomethyl)glycine product concentration such that essentiallyall of the N-(phosphonomethyl)glycine product in the final reactionmixture effluent 45 withdrawn from the second reactor remains dissolved.Optionally, N-(phosphonomethyl)glycine product precipitated in the finalreaction mixture effluent 45 may be separated with the particulatecatalyst as part of the catalyst recycle stream 11. It should berecognized that the temperature of the reaction mixture within thestirred tank reactors 3 and 40 can vary from reactor to reactor. Forexample, since the intermediate aqueous reaction mixture 41 is notfiltered and also contains a lower concentration ofN-(phosphonomethyl)glycine product than does the final reaction mixtureeffluent 45, the temperature of the reaction mixture within the firststirred tank reactor 3 can typically be somewhat lower than thepreferred operating temperature of the reaction mixture in the secondstirred tank reactor 40. Preferably, the first stirred tank reactor 3 isoperated at a temperature of from about 80° C. to about 120° C., morepreferably from about 85° C. to about 110° C. and still even morepreferably from about 95° C. to about 100° C., while the second stirredtank reactor 40 is preferably operated at a temperature of from about80° C. to about 120° C., more preferably from about 85° C. to about 110°C. and even more preferably from about 100° C. to about 105° C.Operating the first stirred tank reactor 3 at a lower temperature isoften advantageous to reduce the rate of formation ofN-methyl-N-(phosphonomethyl)glycine which increases at highertemperatures.

The total pressure in the first and second stirred tank reactors 3 and40 is preferably maintained high enough to prevent the liquid reactionmedium in the oxidation reaction zones from boiling and is generallyfrom about 0 to about 500 psig. Typically, the total pressure in thestirred tank reactors 3 and 40 is from about 30 to about 500 psig. Whenmaintaining the temperature of the reaction mixture in the first andsecond oxidation reaction zones within the preferred temperature rangesdisclosed above, the total pressure maintained within the first andsecond stirred tank reactors 3 and 40 is preferably from about 30 toabout 130 psig and more preferably from about 90 to about 110 psig.

The oxygen partial pressure may vary in different regions of theoxidation reaction zones. Preferably, the oxygen partial pressure in theheadspace above the liquid reaction medium in stirred tank reactors 3and 40 is from about 0.1 to about 35 psia, more preferably from about 1to about 10 psia.

Particularly where the concentration of the N-(phosphonomethyl)iminodiacetic acid substrate in the aqueous feed stream 1 (whichincludes the catalyst recycle stream 11 and any other recycle streamsfrom other parts of the process) is from about 7 to about 12% by weight,and even more particularly is about 9% by weight, it is typicallypreferred for the residence time in the first stirred tank reactor 3 tobe such that the N-(phosphonomethyl)iminodiacetic acid substrateconversion to the N-(phosphonomethyl)glycine product in the firstoxidation reaction zone is at least about 50%, more preferably at leastabout 70%, even more preferably from about 80% to about 95%, still evenmore preferably from about 85% to about 95%, and most preferably about90%. The residence time necessary to achieve the desired degree ofconversion will vary with the oxidation reaction conditions employed inthe first stirred tank reactor 3. Typically, the residence time in thefirst stirred tank reactor 3 is from about 5 to about 50 minutes,preferably from about 10 to about 30 minutes, even more preferably fromabout 14 to about 24 minutes and still even more preferably about 20minutes. The residence time in the second stirred tank reactor 40 istypically from about 1 to about 50 minutes, more preferably from about 1to about 30 minutes, more preferably from about 3 to about 20 minutes,more preferably from about 6 to about 20 minutes, still even morepreferably from about 6 to about 12 minutes and still yet even morepreferably about 8 minutes. The residence time in the first stirred tankreactor 3 is defined relative to the flowrate of the intermediatereaction mixture 41 and the working volume of the reactor. The residencetime in the second stirred tank reactor 40 is defined relative to theflowrate of the final reaction mixture effluent 45 and the workingvolume of the reactor Conversion achieved at a given residence timetends to decrease as the catalyst activity decreases with use, requiringfortification of catalyst activity by reactivation or charging thesystem with fresh catalyst or an increasing the O₂ feed rate.

Preferably, the ratio of the working volume of liquid reaction medium inthe first stirred tank reactor 3 to the working volume of the liquidreaction medium in the second stirred tank reactor 40 is greater than 1,more preferably greater than 1 and up to about 10, even more preferablyfrom about 1.1 to about 5, and still even more preferably from about 1.1to about 2.5.

Normally, when the continuous reactor system comprises two stirred tankreactors in series, the total oxygen feed introduced to the continuousreactor system (i.e., the combined oxygen feed to both stirred tankreactors 3 and 40) and the amount of the total oxygen feed apportionedto each of the stirred tank reactors are adjusted to affect the yieldand quality of the N-(phosphonomethyl)glycine product. In oneembodiment, the total oxygen introduced to the continuous reactor systemper mole of N-(phosphonomethyl)iminodiacetic acid substrate in theaqueous feed stream 1 introduced to the first stirred tank reactor 3 isvaried to control the concentration of N-(phosphonomethyl)iminodiaceticacid substrate in the final reaction mixture effluent 45 withdrawn fromthe second stirred tank reactor 40. The concentration of unreactedN-(phosphonomethyl)iminodiacetic acid substrate in the final reactionmixture 45 is generally minimized to avoid excessive yield losses.Preferably, the concentration of unreactedN-(phosphonomethyl)iminodiacetic acid substrate is no greater than about2000 ppm in the final reaction mixture effluent. However, theconcentration of N-(phosphonomethyl)iminodiacetic acid substrate in thefinal reaction mixture effluent 45 should remain sufficiently high toinhibit the rate at which the N-(phosphonomethyl)glycine productoxidizes to form aminomethylphosphonic acid. The rate ofaminomethylphosphonic acid formation is apparently inverselyproportional to the N-(phosphonomethyl)iminodiacetic acid substrateconcentration. Moreover, it is believed that the presence ofN-(phosphonomethyl)iminodiacetic acid substrate may inhibitover-oxidation of the catalyst and extend the catalyst life.Accordingly, it is preferred that the concentration of theN-(phosphonomethyl)iminodiacetic acid substrate in the final reactionmixture effluent 45 be maintained within a range of from about 200 toabout 2000 ppm, more preferably from about 500 to about 1500 ppm, andmost preferably about 500 to about 700 ppm by weight. Typically, asuitable concentration of the N-(phosphonomethyl)iminodiacetic acidsubstrate in the final reaction mixture 45 is obtained when the totaloxygen introduced to the continuous reactor system is from about 0.5 toabout 5, more preferably from about 1 to about 3, still more preferablyfrom about 1.5 to about 2.5 moles of O₂ per mole ofN-(phosphonomethyl)iminodiacetic acid substrate in the aqueous feedstream 1 introduced to the first stirred tank reactor 3.

In addition, the apportionment of the total oxygen feed to thecontinuous reactor system between stirred tank reactors 3 and 40 isselected to reduce the quantity of by-products in the final reactionmixture effluent 45. The proportion of the total oxygen feed to thecontinuous reactor system introduced into the first stirred tank reactor3 is from about 10% to about 95%, more preferably from about 30% toabout 95%, still more preferably from 50% to about 95% and mostpreferably from about 70% to about 90% with the remaining portion of thetotal oxygen feed being introduced into the second stirred tank reactor40.

In the practice of the present invention, the concentration of unreactedN-(phosphonomethyl)iminodiacetic acid substrate,N-(phosphonomethyl)glycine product and/or oxidation by-products in theintermediate aqueous reaction mixture 41 withdrawn from the firststirred tank reactor 3 and/or in the final reaction mixture effluent 45withdrawn from the second stirred tank reactor 40 may be measured. Basedon these measurements, the total oxygen feed to the continuous reactorsystem and/or the apportionment of the total oxygen feed between thefirst and second stirred tank reactors 3 and 40 may be adjusted tobeneficially affect the yield and quality of theN-(phosphonomethyl)glycine product. The concentration of unreactedN-(phosphonomethyl)iminodiacetic acid substrate,N-(phosphonomethyl)glycine product and/or oxidation by-products can bemeasured using high pressure liquid chromatography (HPLC) or Fouriertransform infrared spectroscopy (FTIR) analysis of stream samples. Inaddition, an in-line FTIR spectrometer may be used to provide real timecompositional analysis of the reactor effluent streams and this dataused in adjusting the oxygen feed practice in the continuous reactorsystem. In-line use of infrared spectroscopy to measure concentrationsof analytes in oxidation reaction mixtures such as those prepared inaccordance with the present invention for use in process control andendpoint detection are described in a U.S. provisional patentapplication No. 60/292,659 entitled “Use of Infrared Spectroscopy forOn-Line Process Control and Endpoint Detection”, filed on May 22, 2001,the entire disclosure of which is expressly incorporated herein byreference.

Normally, when the continuous reactor system comprises two continuousstirred tank reactors 3 and 40 in series, the oxygen feed rate to thefirst reaction zone is preferably from about 0.5 to about 10, morepreferably from about 0.5 to about 5, still more preferably from about1.0 to about 4.0 moles of O₂ per mole ofN-(phosphonomethyl)iminodiacetic acid substrate contained in the aqueousfeed stream 1 introduced into the first reactor 3. The oxygen feed rateinto the second reaction zone is preferably from about 0.5 to about 10,more preferably from about 0.5 to about 5, still more preferably fromabout 2 to about 4 moles of O₂ per mole ofN-(phosphonomethyl)iminodiacetic acid substrate contained in the feedstream to the second reaction zone.

Where the process uses two stirred tank reactors in series, the molarratio of N-(phosphonomethyl)iminodiacetic acid substrate toN-(phosphonomethyl)glycine product in the first reactor is preferablymaintained such that the molar rate of oxidation of theN-(phosphonomethyl)iminodiacetic acid substrate is at least about 10,more preferably at least about 20, even more preferably at least about100, still even more preferably at least about 150, and most preferablyat least about 200 times as fast as the molar rate of oxidation of theN-(phosphonomethyl)glycine product.

Various alternatives to the flow scheme shown in FIG. 3 may be used tocirculate the particulate heterogenous catalyst through the back-mixedoxidation reaction zones within the continuous reactor system. Examplesof such alternative flow schemes are shown in FIGS. 4-6. In each of theflow schemes shown in FIGS. 3-7, the catalyst age may be maintainedwithin a desirable range or controlled near a specific level bycontinuously or intermittently adding fresh catalyst into the catalystrecycle stream(s) or directly into either of the reaction zones. Suchcatalyst age may optionally be further controlled by also continuouslyor intermittently purging a portion of the catalyst from the catalystrecycle stream(s). Often, the amount of purged catalyst is equal to theamount of fresh catalyst added to the system. Intermittent purging andadding of catalyst includes any repeated purging and adding of catalystwhich is not continuous. For example, intermittent purging and addingincludes periodic withdrawal of catalyst from a catalyst recycle stream,with addition of fresh catalyst at a point downstream of the withdrawalpoint within a catalyst recirculation loop. Intermittent purging andadding also includes, for example, withdrawing all the catalyst fromfewer than all the reaction zones at one time, and then adding anentirely fresh batch of catalyst to fewer than all the reaction zones.Intermittent purging and adding further includes, for example,withdrawing all the catalyst from the continuous reactor system at thesame time and then adding an entirely fresh batch of catalyst (e.g.,once the production of N-(phosphonomethyl)glycine product from thereactor system has reached a predetermined target value based on thecalculated useful life of the catalyst load or once the catalystactivity has declined to an extent that economical operation isimpaired). The latter method is typically more preferred. This stemsfrom, for example, the fact that it is often difficult to stabilize thesystem when only portions of the catalyst load are purged and added at agiven time. It is also, for example, difficult to analyze anymodification (e.g., new improvements) to a catalyst without firstremoving all the unmodified catalyst. It should be further noted that atstartup of the continuous oxidation reactor system, it may beadvantageous to operate the system for a time with significantly lessthan the design catalyst loading (e.g., 75% of the design catalystloading) and then to incrementally charge additional catalyst to thesystem to arrive at an optimal catalyst loading at the prevailingoperating conditions.

FIG. 4 shows an embodiment which provides more flexibility by allowingthe catalyst loading into the first and second stirred tank reactors 3and 40 to be manipulated so that a desired greater catalyst loading maybe maintained in the second stirred tank reactor 40 to at leastpartially compensate for the reduced N-(phosphonomethyl)iminodiaceticacid substrate concentration driving force that is typically present dueto the lower substrate concentration in the second reaction zone.Catalyst is continuously or intermittently introduced into the firststirred tank reactor 3 via catalyst feed stream 39. The catalyst is thencontinuously or intermittently withdrawn from the first stirred tankreactor 3 as part of the intermediate aqueous reaction mixture 41,continuously or intermittently introduced into the second stirred tankreactor 40 as part of the intermediate aqueous feed stream 42, andfinally intermittently or continuously withdrawn from the second stirredtank reactor 40 as part of the final aqueous reaction mixture 45. Thecatalyst is then essentially removed from the final aqueous reactionmixture 45 by catalyst filter 9 to form (1) a catalyst recycle stream 11comprising essentially all the catalyst and a residual amount ofN-(phosphonomethyl)glycine product from the final aqueous reactionmixture 45; and (2) a filtrate 13 comprising the bulk ofN-(phosphonomethyl)glycine product from the final aqueous reactionmixture 45. The catalyst recycle stream 11 is divided into the catalystfeed stream 39 and an intermediate catalyst feed stream 50. The catalystfeed stream 39 is recycled back to the first stirred tank reactor 3,while the intermediate catalyst feed stream 50 is recycled back to thesecond stirred tank reactor 40. Preferably, the catalyst is continuouslyor intermittently purged from the continuous reactor system through, forexample, the catalyst purge stream 51 and/or catalyst purge stream 53,and replenished through, for example, catalyst feed stream 55 and/orcatalyst feed stream 57. Catalyst could alternatively or additionally bepurged from catalyst recycle stream 11, and likewise fresh catalystcould alternatively or additionally be added to catalyst recycle stream11 prior to dividing recycle stream 11 into recycle catalyst streams 39and 50. The catalyst may also be at least partially reactivated asdescribed above by intermittently or continuously introducing a reducingagent and/or a supplemental promoter into the continuous reactor system,particularly where the catalyst comprises the deeply reduced catalystdescribed above. The reducing agent and/or supplemental promoter may beintroduced, for example, in the catalyst recycle streams 11, 39 and/or50. Such reactivation may optionally be conducted in one or morecatalyst holding tanks (not shown).

FIG. 5 shows an embodiment wherein each oxidation reaction zone utilizesits own independent particulate catalyst mass. In such an embodiment, anaqueous feed stream 1 comprising the N-(phosphonomethyl)iminodiaceticacid substrate is fed into the first stirred tank reactor 3, wherein itis continuously oxidized in the presence of the first catalyst mass toform an intermediate aqueous reaction mixture 41. This intermediateaqueous reaction mixture 41 is filtered in catalyst filter 9 a toseparate essentially all the first catalyst mass from the intermediateaqueous reaction mixture 41 and form (1) a first catalyst recycle stream11 a comprising essentially all the catalyst from the intermediateaqueous reaction mixture 41; and (2) an intermediate aqueous feed stream60, the filtrate from the filter 9 a, comprising the bulk ofN-(phosphonomethyl)glycine product and unreactedN-(phosphonomethyl)iminodiacetic acid from the intermediate aqueousreaction mixture 41. The first catalyst recycle stream 11 a is fed backinto the first stirred tank reactor 3 via catalyst feed stream 39 a,while the intermediate aqueous feed stream 60 is introduced into thesecond stirred tank reactor 40, wherein further continuous oxidation ofN-(phosphonomethyl)iminodiacetic acid substrate (and C₁ molecules, suchas formaldehyde and formic acid) takes place in the presence of a secondparticulate catalyst mass to form the final reaction mixture effluent45. The final reaction mixture 45, after optional depressurization in aflash tank 17 b, is filtered in catalyst filter 9 b to separate thesecond catalyst mass from the final aqueous reaction mixture 45 and form(1) a catalyst recycle stream 11 b comprising essentially all thecatalyst from the final aqueous reaction mixture 45; and (2) a filtrate13 comprising the bulk of N-(phosphonomethyl)glycine product from thefinal aqueous reaction mixture 45. The catalyst recycle stream 11 b isthen fed back into the second stirred tank reactor 40 via catalyst feedstream 39 b. Preferably, the catalyst mass utilized in the first stirredtank reactor 3 is continuously or intermittently purged through thecatalyst purge stream 20 a, and replenished through catalyst feed stream21 a. Likewise, the catalyst mass utilized in the second stirred tankreactor 40 is preferably continuously or intermittently purged throughthe catalyst purge stream 20 b, and replenished through catalyst feedstream 21 b. The particulate catalyst masses for the first and secondstirred tank reactors 3 and 40 may also be at least partiallyreactivated, as described above, by continuously or intermittentlyintroducing a reducing agent 18 a and 18 b and/or a supplementalpromoter 19 a and 19 b into the respective catalyst holding tanks 5 aand 5 b or at other locations in the continuous reactor system. Forexample, the supplemental promoter may also be added directly to one orboth of the stirred tank reactors 3 and 40.

The catalyst recycle scheme shown in FIG. 5 is advantageous because itprovides flexibility for independently manipulating the catalyst type,age, and loading in each reaction zone. For example, the catalystemployed in the first stirred tank reactor 3 may be tailored to obtainhigh conversion of N-(phosphonomethyl)iminodiacetic acid substrate underthe selected operating conditions in the first oxidation reaction zone,while the catalyst employed in the second stirred tank reactor 40 may beoptimized for improved oxidation of formaldehyde and formic acidby-products and minimal over-oxidation of the N-(phosphonomethyl)glycineproduct. Also, two filter reactor systems, such as the one shown in FIG.5, can tolerate a filter that generates a catalyst recycle stream thatis less concentrated than the desired concentration in a single filterreactor system, such as the system shown in FIG. 3.

In some embodiments, the benefits of a younger catalyst can be greaterin one reaction zone versus another. For example, the effects of anaging catalyst in the first reaction zone (where the bulk ofN-(phosphonomethyl)iminodiacetic acid substrate normally is oxidized)may, in some embodiments, not be as detrimental as the effects of anaging catalyst in the second reaction zone, and the effects of freshcatalyst may likewise be greater in the second reaction zone than in thefirst reaction zone. This may be true, for example, in embodiments wherethe bulk of N-(phosphonomethyl)iminodiacetic acid substrate is oxidizedin the first reaction zone, and the resulting low substrateconcentration in the second reaction zone causes a slower reaction rate.In such an instance, it may be sometimes preferable to use a reactorsystem having a flow scheme like the one shown in FIG. 6. In thisembodiment, catalyst from the particulate catalyst mass utilized in thesecond stirred tank reactor 40 may be continuously or intermittentlypurged from the catalyst recycle stream 11 b via stream 65 andintroduced into the first stirred tank reactor 3 via the catalystrecycle stream 11 a, thereby extending the useful life of the catalystin the overall process. Such a scheme is particularly advantageous wherethe catalyst comprises a costly material, such as a noble metal.Normally, in this embodiment, fresh catalyst is introduced only into thesecond reaction zone via catalyst feed stream 21 b, while catalyst ispurged from the process only from the first reaction zone via catalystpurge stream 20 a. The average catalyst age (i.e., cumulative time thatthe catalyst has been used to catalyze the oxidation reaction) in thesecond stirred tank reactor 40 is preferably from about 20 to about 65%of the average age of the catalyst utilized in the first stirred tankreactor 3. The average amount of N-(phosphonomethyl)glycine productproduced per pound of catalyst in the second stirred tank reactor 40preferably is from about 5 to about 30% the average amount ofN-(phosphonomethyl)glycine product produced per pound of catalyst in thefirst stirred tank reactor 3.

It should be recognized that in some embodiments, it is more preferableto recycle the catalyst in the opposite direction as that shown in FIG.6 (i.e., the catalyst flows co-currently with the substrate). In thoseinstances, fresh catalyst is continuously or intermittently introducedinto the first reaction zone, catalyst from the particulate catalystmass utilized in the first stirred tank reactor 40 is continuously orintermittently purged from the catalyst recycle stream 11 a andcontinuously or intermittently transferred to the second reaction zoneand catalyst in the second reaction zone is continuously orintermittently purged from the reactor system. In such an embodiment,the average catalyst age in the first stirred tank reactor 3 ispreferably from about 33 to about 80% of the average age of the catalystin the second stirred tank reactor 40. The average amount ofN-(phosphonomethyl)glycine product produced per pound of catalystconsumed in the first stirred tank reactor 3 preferably is from about 75to about 90% the average amount of N-(phosphonomethyl)glycine productproduced per pound of catalyst consumed in the second stirred tankreactor 40.

In the embodiments shown in FIGS. 5 and 6, either external heat transferrecirculation loop 15 or 47 may also be a catalyst recycle loop in thesame manner as shown in FIG. 2B, rather than being independent of thecatalyst recycle streams 11 a or 11 b, respectively. For such a combinedloop, the catalyst filters 9 a and 9 b are preferably continuouscross-flow filters.

In processes including operating two oxidation reaction zones in series,particularly two stirred tank reactors 3 and 40 in series, it isdesirable to achieve a high rate of mass transfer in the first oxidationreaction zone. Therefore, it is preferred to introduce the O₂-containinggas, preferably a gas containing at least about 95 mole % O₂, typicallyabout 98 mole % O₂, directly into the reaction mixture in the firststirred tank reactor 3 through a sparger located just below or near theimpeller and also to minimize the back-mixing of gases to maximize theoxygen concentration driving force for high mass transfer in the firstoxidation reaction zone. For equivalent pressures and oxygen conversion,the average oxygen spatial concentration is expected to be higher inreaction environments with minimal gas phase back-mixing. Near thesparger, for example, the oxygen partial pressure in the undissolvedgases is normally greater than in other regions in the reactor, such asnear the interface between the liquid reaction medium and the headspace.However, in the second reaction zone, where theN-(phosphonomethyl)iminodiacetic acid substrate concentration istypically much lower, mass transfer demands and the need for a highoxygen concentration driving force are considerably less. Thus,back-mixing of gases is more easily tolerated in the second oxidationreaction zone and, in some instances, is preferred. The deeply reducednoble-metal-on-carbon catalyst preferred in the practice of the presentinvention is more susceptible to over-oxidation in reaction environmentshaving pockets of high oxygen partial pressures in the undissolvedgases, especially at low concentrations ofN-(phosphonomethyl)iminodiacetic acid substrate such as thoseencountered in the second oxidation reaction zone. By back-mixing thegas phase in the liquid reaction medium within the second reaction zone,average oxygen spatial concentration is decreased and the stability ofsuch a catalyst is enhanced.

Various reactor modifications may be employed to maintain a more uniformlow oxygen partial pressure in the undissolved gases in the reactionmixture contained in the second reaction zone. One preferred alternativeis to select an impeller system for the second stirred tank reactor 40that is adapted to provide a high rate of gas induction from theheadspace interface into the reaction mixture such as A340 up-pumpingaxial flow impeller system available from Lightnin (Rochester, N.Y.,U.S.A.). Such an impeller system draws gas from the headspace into theliquid reaction mixture so that the difference between the oxygenpartial pressure of the gas being drawn into the liquid reacting mediumand the oxygen partial pressure of the headspace gas is reduced, therebylowering the average oxygen spatial concentration in the undissolvedgases in the reaction mixture. In addition, the second stirred tankreactor 40 may be modified so that the O₂-containing gas is fed into theheadspace above the reaction mixture rather than being sparged directlyinto the liquid reaction mixture. This will even further reduce theoccurrence of pockets of high oxygen concentration. Alternatively, theaverage oxygen spatial concentration may be reduced, by introducing theheadspace gas within the second stirred tank reactor 40 into the liquidreaction mixture through the impeller. A commercially available exampleof such an impeller system including a hollow shaft for gas transport isthe DISPERSIMAX system, sold by Autoclave France (Nogent-sur-Oise Cedex,France). Another possibility is to decrease the O₂ concentration in theO₂-containing gas introduced into the second stirred tank reactor 40(e.g., air may be used as the oxygen source supplied to the secondoxidation reaction zone).

In a further modification, the second continuous stirred tank reactor 40is replaced by an ejector nozzle loop reactor. A schematic diagram ofsuch a reactor is shown in FIG. 7. Here, an aqueous feed stream 901comprising at least a portion of the intermediate aqueous reactionmixture 41 withdrawn from the first oxidation reaction zone is pumpedinto inlet 903 and ejected through a nozzle 907 into a mixing chamber909 into which an O₂-containing gas is also introduced via inlet 911(i.e., the O₂-containing gas is introduced into the throat of theventuri nozzle 907). This creates a high mass transfer coefficient foroxygen transfer into the aqueous feed 901. Because of this high oxygenmass transfer coefficient and the high agitation within the reactorvessel 913 caused by the nozzle 907, the average oxygen spatialconcentration in the undissolved gases in the liquid reaction mixture915 is low. The reaction mixture effluent 917 is withdrawn from anoutlet 919 near the bottom of the reactor vessel 913, cooled in a heatexchanger 921, and filtered by catalyst filter 922, preferably across-flow filter. Catalyst separated from the reaction mixture effluent917 is recirculated back to the reactor 913 via catalyst recycle stream923 using a pump 925. The filtrate 927, which contains the bulk of theN-(phosphonomethyl)glycine product, is forwarded to be purified and/orconcentrated in additional steps. Operation and design of ejector nozzleloop reactors is described by van Dierendonck, et al. in “Loop VenturiReactor-A Feasible Alternative to Stirred Tank Reactors?”, Ind. Eng.Chem. Res. 37, 734-738 (1998), the entire disclosure of which isincorporated herein by reference. A commercially available example of anejector nozzle loop reactor is the BUSS loop reactor sold by KvaernerBuss CPS (Pratteln, Switzerland). It should be understood that inaddition to providing a second or subsequent oxidation reaction zone ina continuous reactor system comprising multiple oxidation reaction zonesin series, an ejector nozzle loop reactor could likewise suitablyprovide the first oxidation reaction zone. The oxidation reactionconditions and operating parameters for an ejector nozzle loop reactorare similar to those described above for oxidation reaction zonesprovided by stirred tank reactors.

Much of the preceding discussion has focused on continuous reactorsystems utilizing a heterogeneous particulate catalyst slurry andcomprising at least two stirred tank reactors in series providingoxidation reaction zones substantially back-mixed in at least the liquidphase. However, it should be recognized that reactor configurationsother than stirred tank reactors may be equally or more suitable thanstirred tank reactors for one or more of the oxidation reaction zones orcould be used in combination with multiple stirred tank reactor stages.Furthermore, many such alternative reactor configurations are likewisesuitable for use in continuous reactor systems including a singleoxidation reaction zone. One of the disadvantages of a continuousreactor system including one or more stirred tank reactors utilizing aparticulate catalyst slurry is the capital and operating cost associatedwith a catalyst recycle mechanism including a catalyst filter or othercatalyst separation means necessary to recover theN-(phosphonomethyl)glycine product. Accordingly, reactor configurationsin which the catalyst can remain in the oxidation reaction zone mayprovide an economic advantage in some applications. Two examples of suchreactor configurations are fixed catalyst bed reactors and fluidized bedreactors. A further advantage of fixed bed reactors and fluidized bedreactors is that they can be operated in a manner to exhibit plug flowcharacteristics which tends to produce lower concentrations ofundesirable byproducts (e.g., N-methyl-N-(phosphonomethyl)glycine), and,consequently, a greater N-(phosphonomethyl)glycine product yield.

FIG. 8 shows an example of a fixed bed reactor 500 in accord with oneembodiment of the present invention. Disposed within the reactor 500 isa primary oxidation reaction zone comprising a primary fixed bed 501containing an oxidation catalyst, preferably the deeply reduced catalystdescribed above. A fixed bed support 502 is preferably positioned withinthe reactor 500 to provide an upper chamber 503 and a lower chamber 504above and below the fixed bed 501, respectively. An aqueous feed stream505 comprising the N-(phosphonomethyl)iminodiacetic acid substrate iscontinuously or intermittently introduced into the upper chamber 503 anddistributed over the fixed bed 501 by spray nozzles 506 or otherconventional liquid distribution system. An O₂-containing gas islikewise introduced into the upper chamber 503. As the O₂-containing gasflows cocurrently through the fixed bed 501 with the descending flow ofliquid reaction mixture, the N-(phosphonomethyl)iminodiacetic acidsubstrate is continuously oxidized. A primary reactor effluent 507comprising N-(phosphonomethyl)glycine product is withdrawn from thelower chamber 504 along with a vapor stream comprising CO₂.

Although downward, cocurrent flow of the liquid reaction mixture and theO₂-containing gas through the fixed bed 501 is shown in FIG. 8, itshould be understood that various flow combinations are possible. Forexample, the aqueous feed stream 505 and the O₂-containing gas could beintroduced into the lower chamber 504 of the reactor 500 and flowcocurrently, upward through the fixed bed 501. Alternatively, the liquidreaction mixture and the O₂-containing gas can flow countercurrentlythrough the fixed bed 501, the O₂-containing gas being introduced intothe lower chamber 504 and the aqueous feed stream 505 being introducedinto the upper chamber 503 or vice versa.

The temperature within the oxidation reaction zone of the fixed bedreactor 500 is preferably in the range of from about 20° C. to about180° C., more preferably from about 60° C. to about 140° C., still morepreferably from about 80° C. to about 130° C., and yet still morepreferably from about 90° C. to about 120° C. Although the reactionsystem may optionally be operated adiabatically, adverse effects on thecatalyst or undue formation of by-products may result from excessivetemperatures encountered in adiabatic operation within a primaryoxidation reaction zone (i.e., a reaction zone into which a substantialfraction of unconverted substrate is introduced). Where the substrate isof limited solubility (e.g., N-(phosphonomethyl)iminodiacetic acid), atemperature of at least the saturation temperature is preferablymaintained at the reactor inlet in order to prevent substrate solidsfrom being deposited in the bed. However, effect on by-product formationand catalyst deterioration require that the maximum temperature bemaintained within the ranges outlined above. As a practical matter, thislimits the extent of conversion of substrate that may be achieved in anadiabatic fixed bed to not greater than about 10% by weight on a totalreaction mixture basis, preferably not greater than about 7%, moretypically in the range of about 3% to about 5%. Where the substrate is asalt, the conversion is not constrained by solubility of substrate, butis still limited in the aforementioned range by effects on catalyst andby-product formation.

To achieve a more substantial conversion in a single fixed bed,exothermic reaction heat must be removed from the reaction system.Although the reaction zone as such may be operated adiabatically, heatmust be removed from somewhere in the reaction system so that thedifference in unit weight sensible heat content between the reactionmixture and the aqueous feed stream is maintained at a value less thanthe exothermic reaction heat generated in the reaction zone per unitweight of the aqueous feed stream. As described below, measures toremove reaction heat may include cooling the reaction zone, orintroducing a cooled recirculation stream with the aqueous feed mixture.By applying cooling in such manner, the conversion expressed as thedifference between reaction mixture composition and feed composition maybe increased to above 10% or even above 15%. Where the substrate andproducts are water-soluble salts, the conversion can be increased to20%, 30%, or even 50%.

Controlling the temperature within the oxidation reaction zone of afixed bed reactor is typically more difficult as compared to temperaturecontrol in a back-mixed reactor system. As shown in FIG. 8, primaryreactor effluent 507 may be divided into a primary product fraction 508and a recirculation fraction 509 which is cooled externally of thereaction zone and returned to the inlet of reactor. Typically at leastabout 5%, preferably at least about 33%, more preferably from about 50%to about 90%, and even more preferably from about 60% to about 80% ofthe primary reactor effluent 507 exiting the reactor is diverted torecirculation fraction. Expressed another way, the ratio of thevolumetric flow rate of the recirculation fraction 509 to the volumetricflow rate of the primary reaction product fraction 508 is typically atleast about 0.05:1, preferably at least about 0.5:1, more preferablyfrom about 1:1 to about 10:1, and most preferably from about 1.5:1 toabout 5:1. The recirculation fraction is cooled externally of the fixedbed before it is returned to the reactor, cooling being effected in aheat exchanger 510. In one embodiment shown in FIG. 8, the primaryreaction mixture exiting the reactor is divided into a primary productfraction and recirculation fraction and the product fraction removedbefore the recirculation fraction is passed through the heat exchanger.This alternative may be advantageous in certain embodiments, forexample, where the primary product fraction contains unreactedN-(phosphonomethyl)iminodiacetic acid substrate or by-product C₁compounds which are to be oxidized in a further reaction zone. Inaddition this flow scheme may be advantageous if the primary productfraction is evaporated (e.g., to concentrate and recover theN-(phosphonomethyl)glycine product). In an alternative shown in FIG. 8,it may be preferable to produce a cooled product fraction by passing theentire primary reaction mixture, or substantially the entire primaryreaction mixture through an external heat exchanger such as thatillustrated at 510, and thereafter dividing the cooled primary reactionmixture stream into a recirculation stream and a cooled primaryfraction. Throughout the Figures, dashed lines with directionalarrowheads indicate optional, alternative or additional streams.

The cooled recirculation fraction 509 and the aqueous feed stream 505are mixed to produce a combined inlet stream for the primary reactionzone. Due to the reaction, the recirculation fraction is relativelydepleted in N-(phosphonomethyl)iminodiacetic acid substrate, a factorwhich can be exploited to maximize productivity by introducing anaqueous feed mixture having a high substrate content, includingsubstrate concentrations in excess of the solubility limit of thesubstrate in the aqueous phase of the feed mixture. Because therecirculation fraction is relatively depleted in substrate, the effectof mixing is to produce a combined inlet stream which has asignificantly lower substrate content than the aqueous feed mixture.This dilution effect allows the feed mixture to be much moreconcentrated than would otherwise be possible. For example, the feedmixture may comprise a slurry of N-(phosphonomethyl)iminodiacetic acidin a saturated, or substantially saturated aqueous solution thereof,which might otherwise tend to cause plugging of the fixed catalyst bed.Mixing with the recirculation fraction reduces theN-(phosphonomethyl)iminodiacetic acid content sufficiently to dissolvethe slurry solids and provide a combined inlet stream that issubstantially free of solid substrate. Typically, heat from therecirculation fraction also causes the temperature of the combinedstream to exceed that of the aqueous feed mixture, further contributingto dissolution of the substrate solids. Moreover, because the oxidationof N-(phosphonomethyl)iminodiacetic acid to N-(phosphonomethyl)glycineis substantially zero order until a high conversion has been attained,the dilution effect does not adversely affect the reaction rate.Optionally, the aqueous feed mixture and recirculation fraction can bedirected to a mixing tank to assure that solids have been dissolvedbefore the combined inlet stream is introduced into the fixed catalystbed. In this manner it is feasible to introduce an aqueous feed streamcomprising between about 8% and about 15% by weightN-(phosphonomethyl)iminodiacetic acid, and to produce a combined inletstream by mixing this aqueous feed stream with a primary reactorrecirculation fraction comprising between about 0.5% and about 5% byweight N-(phosphonomethyl)glycine. As described hereinbelow,significantly higher concentrations can be processed where the substrateis a water-soluble salt of N-(phosphonomethyl)iminodiacetic acid and theproduct is a water-soluble salt of N-(phosphonomethyl)glycine.

It should be understood that, in reducing the substrate content of thecombined inlet stream, dilution of the aqueous feed stream with therecirculation stream further serves to reduce the difference insubstrate content between the combined catalyst bed liquid phase inletstream and the catalyst bed liquid exit stream, allowing this differenceto be maintained in the ranges which can be tolerated in an adiabaticreaction zone, as described above. Maintaining such limitation onproportional conversion of substrate within the reaction zone remainsimportant in the system of FIG. 8 inasmuch as the catalyst bed itselfmay still be operated substantially adiabatically, though the overallreaction system, including the recirculation loop, is not.

In an alternative embodiment of the invention (not shown), a continuousreactor system may comprise a second oxidation reaction zone into whichpart or all of the primary reaction product fraction 508 may becontinuously introduced for further conversion ofN-(phosphonomethyl)iminodiacetic acid substrate and oxidation of C₁by-products. In such an embodiment, all of the primary reactor effluent507 may be diverted into the heat exchanger recycle loop and the primaryproduct fraction 508 removed from the recirculation fraction 509downstream of the heat exchanger 510. In this manner, some of theexothermic heat of reaction would be removed from the primary productfraction 508 before introduction into the second oxidation reactionzone. The second reaction zone contains an oxidation catalyst and may beback-mixed, as provided within a continuous stirred tank reactor, or maycomprise a second fixed catalyst bed. ResidualN-(phosphonomethyl)iminodiacetic acid substrate in the primary reactionproduct fraction is continuously oxidized to N-(phosphonomethyl)glycineproduct in the second reaction zone. In a preferred embodimentcomprising two or more fixed bed reactors in series, the reaction iscarried to a high conversion of N-(phosphonomethyl)iminodiacetic acidsubstrate to N-(phosphonomethyl)glycine product in the primary oxidationreaction zone (e.g., at least about 95%, preferably at least about 98%)which is quite feasible because the oxidation proceeds as an apparentzero order reaction until only a small fraction of the originalN-(phosphonomethyl)iminodiacetic acid substrate, e.g., as low as about0.2 ppm or lower, remains in the liquid phase. By operation in thismanner, the heat load of the reaction is very predominantly dissipatedvia the heat exchanger 510 in the recirculation loop of the primaryreactor, and the second oxidation reaction zone can be operatedessentially adiabatically with only modest temperature increase.Optionally, the reaction heat can be removed to a cooling fluid in aninternal heat exchanger (e.g., cooling coils) positioned within thesecond oxidation reaction zone. Such an arrangement withoutrecirculation allows the liquid phase reaction mixture to be passedthrough the second fixed bed without undermining the plug flowcharacteristic (i.e., substantially without back-mixing of the liquidphase via recycle). Plug flow operation is desirable in the secondreaction zone since the oxidation of substrate toN-(phosphonomethyl)glycine product becomes essentially first order athigh conversions. Plug flow operation maximizes the kinetic drivingforce for extinguishing the residual N-(phosphonomethyl)iminodiaceticacid substrate and reduces the amount of by-products fromover-oxidation.

Preferably, both the primary and second catalyst beds contain a noblemetal on carbon catalyst which is effective for oxidation of both theN-(phosphonomethyl)iminodiacetic acid substrate and the C₁ by-products,formaldehyde and formic acid. Since the noble metal functions primarilyto catalyze the oxidation of the C₁ by-products, while the oxidation ofthe N-(phosphonomethyl)iminodiacetic acid substrate is primarilycatalyzed by the carbon, an alternative embodiment of the inventioncomprises the use of a primary fixed bed that consists essentially ofcarbon catalyst, or has a significantly lower noble metal content thanthe catalyst deployed in a second fixed bed. The second reactorcomprises a noble metal on carbon catalyst to assure oxidation of the C₁by-products. Inasmuch as the C₁ oxidation is substantially first orderin any case, it proceeds more effectively under the essentially plugflow conditions that are conveniently maintained in the second reactor.The heat load may be sufficiently modest in the second reactor that itcan be operated without an external heat exchanger and without backmixing or recirculation. In a fixed bed system, it may be feasible touse a catalyst having a lower noble metal loading per unit weight ofcatalyst than may be optimal for a continuous back-mixed reactionsystem.

In a still further embodiment, a third fixed bed reaction zone can beprovided, which also preferably comprises a fixed bed containing a noblemetal on carbon catalyst, and which can be operated substantially inplug flow and optionally, in fact preferably, under adiabaticconditions. This option may be of particular value where the firstreactor uses only a carbon catalyst. Thus, the oxidation ofN-(phosphonomethyl)iminodiacetic acid substrate toN-(phosphonomethyl)glycine product proceeds well in the presence of thecarbon catalyst in the primary oxidation reaction zone, but C₁by-products tend to accumulate in the primary reactor effluent. Thesecond reaction zone promotes both extinction of the substrate andoxidation of C₁ by-products, both of which are essentially first orderreactions that are promoted by the essentially plug flow operation ofthe second fixed bed. Residual C₁ compounds are effectively extinguishedin the third fixed bed oxidation reaction zone.

In a still further embodiment, the continuous reactor system maycomprise a plurality of shorter (or shallower) fixed bed reactors inseries such that the intermediate reaction mixture effluent exiting onestage is passed through the following stage. This embodiment varies fromthe two or three reactor system described above in that only modestconversion would be achieved in any of the series of relatively shallowfixed bed reactor stages. Since substrate conversion in any one bed isrelatively limited, each bed can be operated substantially adiabaticallywith a heat exchanger being placed between each successive shorter fixedbed reactor and the immediately succeeding reactor to cool the reactionmixture so that the temperatures of the reaction mixture do not exceedthe desired operating temperature in any of the fixed bed reactors.Where the series comprises more than two reactors, it may be necessaryto cool only the intermediate reaction mixture exiting the first one,two or three fixed bed reaction zones, after which it may be feasiblefor the remainder to operate adiabatically. The reaction temperature ina fixed bed reactor may also be controlled by, for example,incorporating separate channels or conduits within the fixed bed throughwhich a cooling medium may be passed. It may be noted that, in thisembodiment, not all the reactors of the series need necessarily be fixedbed reactors. For example, the first reactor of the series could be acontinuous stirred tank reactor within which a substantially zero orderoxidation of N-(phosphonomethyl)iminodiacetic acid substrate toN-(phosphonomethyl)glycine product may be carried to a substantialconversion to produce an intermediate reaction mixture, and theintermediate reaction mixture may be transferred to a fixed bed reactor,or a series of fixed bed reactors for completing the conversion andoxidizing residual C₁ by-products.

As described above, catalyst utilized in fixed bed reactors may take avariety of forms and comprise different types of supports including bothpellet supports and monolithic supports. As shown in FIG. 8, it isgenerally preferred that the oxidation catalyst contained within thefixed bed 501 be in the form of a pellet (e.g., the deeply reducedcatalyst described above comprising a carbon pellet support having anoble metal deposited thereon). Such pellet catalysts typically have aparticle size of from about 1 mm to about 10 mm, more preferably fromabout 1.5 mm to 5 mm. A typical packing density for noble metal oncarbon catalyst particles in this preferred size range is from about 0.1to about 1.4 g/l, more preferably 0.25 to about 0.6 g/l, even morepreferably from about 0.3 to about 0.4 g/ml. It has further beendetermined that the noble metal loading in a noble metal on carboncatalyst that is used in a fixed bed may be lower relative to theloading on a comparable catalyst for use in a slurry reactor. Forexample, effective operation of a fixed bed has been demonstrated usingonly a 2 wt. % Pt on carbon catalyst within the 200 to 2000 m²/cm³ BETsurface to liquid holdup volume ratio noted below. Generally, a loadinglower than 3 wt. % Pt may be satisfactory. Where the catalyst comprisesplatinum on carbon, the platinum loading on the catalyst may be lessthan 70% of the loading required in a slurry reactor.

Another advantage of a fixed bed reactor is that by combining differentcatalysts, the catalyst activity can be selectively varied over thelength of fixed bed reactor stage or from one stage to the next in thedirection of reaction mixture flow. For example, a less active catalyst(e.g., a carbon only catalyst) may be deployed in the upstream portionof a fixed bed reactor stage or in earlier stages of a multi-stage fixedbed reactor system and a more active catalyst (e.g., a deeply reducednoble metal on carbon catalyst) may be deployed in the downstreamportion of a fixed bed reactor stage or in later stages of a multi-stagesystem. Alternatively, the fixed bed may comprise a combination ofoxidation catalyst bodies and other means for promoting gas/liquid masstransfer such as rings, saddles, or structured packing. The rings,saddles, or other inert packing functions as a diluent for the catalyst,thereby modulating the activity of the catalyst bed. In this manner, theactivity of the catalyst bed may be varied in the direction of fluidflow as a function of variation of the surface area of the catalystbodies relative to the surface area of the inert packing. Such avariation in catalyst activity acts to offset the decliningconcentration of the N-(phosphonomethyl)iminodiacetic acid substrate inthe reaction mixture while reducing catalyst costs and noble metallosses from the process.

The tendency of fixed bed reactors to produce lower concentrations ofundesirable byproducts due to plug flow characteristics may be enhancedby using a ratio of effective catalyst surface area to liquid in theworking volume which is significantly greater than the ratio used intypical back-mixed (i.e., well-mixed) reactors. In fact, the need tocool the reaction mixture to reduce impurity formation may be reduced orentirely eliminated by using such a ratio. This is due to the fact thatthe large effective catalyst surface area increases the reaction rateand consequently reduces the liquid residence time. The reducedresidence time, in turn, tends to reduce the formation of impuritiesthat are formed by homogeneous reactions, particularlyN-methyl-N-(phosphonomethyl)glycine. In this embodiment, the ratio ofcatalyst BET surface area to volume of liquid (liquid holdup) in theworking volume of the fixed bed reactor preferably is at least about 3m²/cm³, more preferably from about 100 to about 6000 m²/cm³, and evenmore preferably from about 200 to about 2000 m²/cm³. In someapplications a catalyst BET surface area to liquid holdup in the reactormay most preferably be in the range of about 400 to about 1500 m²/cm³.The volumetric ratio of liquid holdup to total bed volume in the fixedbed is preferably in the range of between about 0.1 and about 0.7. Incertain embodiments the low liquid residence time and high surface tovolume ratio may make it advantageous to operate a fixed bed reactor ina relatively high temperature range of, e.g. 150° C., wherein theintegrated average temperature of the liquid phase across the liquidphase flow path in the primary fixed bed is between about 80° C. andabout 130° C., preferably 105° C. to 120° C.

Fixed bed reactors containing a catalyst in monolith form (e.g.,comprising a honeycomb support such as that shown in FIG. 1) aresometimes more preferred than reactors containing a fixed bed ofdiscrete catalyst particles. This is due to the fact that a fixed bed ofcatalyst particles may be subject to clogging if theN-(phosphonomethyl)iminodiacetic acid substrate contained in the aqueousfeed stream 505 precipitates to any significant degree in the oxidationreaction zone. Accordingly, it is typically required that theconcentration of the N-(phosphonomethyl)iminodiacetic acid substrate inthe aqueous feed stream 505 not exceed the saturation concentration atthe reactor feed temperature, which may significantly limit throughput.However, if the fixed bed 501 comprises a catalyst in the form of ahoneycomb or similar monolith, the channels therein can be madesubstantially straight and with a large enough cross-section so thatthey will not be clogged by a reaction mixture containing a slurry ofsolid N-(phosphonomethyl)iminodiacetic acid substrate. Even if a packedbed reactor is not subject to plugging, the monolith can be operatedwith substantially lower pressure drop. This potential advantage ofutilizing a monolithic catalyst in a fixed bed reactor must be weighedagainst the increased costs associated with production of the monolithsupports as compared to the often significantly cheaper pellet orparticulate supports generally preferred in the practice of the presentinvention. This is particularly true where multiple fixed bed stages areemployed with separate N-(phosphonomethyl)iminodiacetic acid substratefeed streams to each stage, thereby avoiding the need for a highsubstrate concentration in the feed stream to the first fixed bed stageto obtain the desired throughput.

The amount and pressure of oxygen within the oxidation reaction zone mayvery significantly depending upon a variety of considerations, includingtotal pressure, bed temperature profile, superficial gas velocity,catalyst bed volume, specific catalyst type and geometry, feedconcentrations, number of oxidation reaction zones, productivity as wellas other factors.

To achieve high conversion in a packed bed reactor system, each of thefixed bed reaction zones, especially the primary reaction zone, ispreferably operated under relatively high oxygen partial pressure topromote oxygen transfer to the liquid phase. Preferably, the integratedaverage oxygen partial pressure over the liquid phase flow path in theprimary oxidation reaction zone is at least about 50 psia, morepreferably at least about 100 psia, even more preferably at least about200 psia. In some embodiments, integrated average oxygen partialpressures in the range of about 300 psia to about 500 psia may beappropriate. Oxygen content of the gas phase at the gas exit of thereactor may be in the range of 20% to 30% or even lower. Oxygen transfermay also be promoted by the high ratio of catalyst surface area tovolume ratio of liquid phase reacting mixture in a fixed bed reactor asdescribed above. Oxygen utilization in the primary reaction zone ispreferably between about 50% and about 95%. Typically, oxygen is fed tothe reactor in a quantity of from about 1.5 to about 10 mole O₂/moleN-(phosphonomethyl)iminodiacetic acid substrate introduced to thereactor.

The total operating pressure in the fixed bed reactor 500 may typicallybe higher than that in a stirred tank reactor and is preferably fromabout 0 to about 1000 psig, more preferably from about 300 to about 1000psig, and even more preferably from about 100 to about 300 psig.

Generally, a somewhat lower oxygen partial pressure may be preferred ina second and/or third fixed bed oxidation reaction zone(s) in order toavoid over-oxidation of the catalyst and compromising its effectivenessin oxidation of C₁ by-products. Thus, in the second or third reactionzone, the integrated average oxygen partial pressure along the liquidflow path is preferably between about 30 psia and about 300 psia, morepreferably between about 30 psia and about 100 psia. Alternatively, theprimary fixed bed reactor of FIG. 8 and/or the second or third fixed bedreactor(s) in the series may be operated using an oxidant other thanmolecular oxygen, for example, H₂O₂, in which case the total reactionpressure and partial pressure of oxygen may be substantially lower thanas described above.

To protect the catalyst against over-oxidation, it is generallypreferred that oxygen partial pressure at the liquid exit of any fixedbed reactor be not greater than about 100 psia, and is preferablybetween about 10 psia and about 50 psia. It is also preferred that theoxygen partial pressure not exceed about 50 psia at any location in thefixed bed wherein the N-(phosphonomethyl)iminodiacetic acid substratecontent of the liquid phase is less than 0.2 ppm; more preferably, theoxygen partial pressure is maintained below about 50 psia at anylocation in the bed wherein the substrate content of the liquid phase isless than greater than about 0.1 ppm.

By observing the operational guidelines set forth above and particularlyin connection with fixed bed reactors comprising a deeply reduced noblemetal on carbon catalyst as described above,N-(phosphonomethyl)iminodiacetic acid substrate may be converted toN-(phosphonomethyl)glycine product in a single fixed bed reactor atproductivities on a reactor volume basis in the range of from about 0.05to about 4 gmole/1/hr, more typically from about 0.2 to about 2gmole/1/hr.

In accordance with the invention, a fixed bed reactor may be operated ata substantial throughput, provided that adequate heat transfer capacityis provided, as depicted, for example, in FIG. 8. Generally, therelative liquid feed rate to the reactor and the reactor volume are suchthat a fixed bed reactor may be operated at a liquid hourly spacevelocity between about 0.5 hr⁻¹ and about 20 hr⁻¹, as computed on thebasis of total catalyst bed volume, at N-(phosphonomethyl)iminodiaceticacid substrate conversions in excess of about 50%. Higher conversions,in excess of 95% or 98% can be achieved at liquid hourly spacevelocities in the range of from about 0.5 to about 5 hr⁻¹. It will beunderstood that the liquid hourly space velocity is based on the totalliquid phase feed stream. Thus, in the reaction system depicted in FIG.8, the liquid phase feed stream comprises the combined inlet streamproduced by mixing the aqueous feed mixture stream and the recirculationstream, as well as any other recycle or cross-flow streams that might beintroduced into a fixed bed reactor in accordance with a particularprocess flow sheet. Typically, the fixed be is cylindrical and ofcircular cross-section and the height to diameter ratio is selected soas to provide adequate liquid distribution over the bed and appropriategas superficial velocity for mass transfer characteristics. Typically,the height to diameter ration for a fixed bed used in the continuousoxidation reactor systems of the present invention is greater than one,more preferably from about 3 to about 40.

It is difficult to maintain a constant catalyst activity and selectivityover time in a fixed bed reactor. Eventually the activity andselectivity of the catalyst decreases to an unacceptable level such thatthe reactor system may have to be shut down to allow replacement and/orreactivation of the catalyst. This is a disadvantage as compared to thecontinuous reactor systems including one or more stirred tank reactorsdescribed above where catalyst replacement and/or reactivation can occurwhile the reactor system remains online. The problem of catalyst removaland reclamation can be resolved by providing duplicate fixed bedreactors that are valved in parallel to the remainder of the reactionsystem, and operating them on an alternating basis. Catalyst can beremoved from the reactor that is out of service and replaced with freshcatalyst; or catalyst reactivation can be conducted in situ in thereactor that is off line.

In accordance with another embodiment of the present invention, theoxidation of the N-(phosphonomethyl)iminodiacetic acid substrate iscarried out in a circulating fluidized bed reactor utilizing aparticulate heterogenous catalyst, preferably the deeply reducedparticulate catalyst described above. A circulating fluidized bedreactor typically provides a greater mass transfer coefficient than astirred tank reactor and may be operated in a manner to substantiallyretain the particulate catalyst within the oxidation reaction zone suchthat a catalyst filter or other catalyst separation means may not berequired at all or at least substantially reduced in size and pressuredrop requirements. FIG. 9 shows an example of a circulating fluidizedbed reactor 400 defining an oxidation reaction zone therein. An aqueousfeed stream 401 comprising the N-(phosphonomethyl)iminodiacetic acidsubstrate is pumped into the top of the reactor 400 through an inlet 403of draft tube 405 and discharged near the bottom of the reactor 400 intoliquid reaction medium 406 in contact with the catalyst particles 407.An O₂-containing gas may be sparged into the reaction mixture through anozzle 409 at the bottom of the reactor 400. Reaction solution 412 iswithdrawn from the reaction zone at an overflow port 411 and vaporcomprising CO₂ is vented through the top of the reactor. The reactor hasa reacting mixture circulation withdrawal port 413 located well abovethe discharge mouth of draft tube 405 but below the overflow port 411.Reacting mixture having particulate catalyst suspended therein iswithdrawn at port 413, circulated via an external loop 420 through aheat exchanger 421 for removal of reaction heat, and then combined withfeed stream 401 for reintroduction into the reactor via draft tube 405.By maintaining a high rate of circulation in loop 420 relative to therate of supply of feed 401 and withdrawal of reaction solution 412, anupward flow velocity is established in a lower slurry region of theoxidation reaction zone generally below port 413 that is much higherthan the upward flow velocity in a upper decantation region of theoxidation reaction zone generally above port 413. The equipment is sizedand the recirculation flow controlled so that upward velocity in thelower slurry region is well above the sedimentation velocity of thecatalyst particles 407 and therefore effective to maintain the catalystin suspension (i.e., entrained) in the reaction medium within the slurryregion. However, the upward velocity in the decantation region aboveport 413 is well below the sedimentation velocity of the catalystparticles 407, allowing separation of a relatively clear reactionsolution decantate 412 which exits through port 411. Typically, the sizeof the catalyst particles 407 utilized in a reactor such as that shownin FIG. 9 is from about 200 μm to about 1000 μm. Smaller catalystparticles that might be entrained in the decantate 412, for example atstartup, may be removed with a polish filter (not shown). Thus, theslurry catalyst is maintained within the reactor obviating the need forfiltration, or at least for a filter having the capacity to removecatalyst at the rate that would be required if the concentration ofcatalyst in the forward flowing reaction solution 412 were comparable tothe concentration of catalyst in the slurry region.

For removal of catalyst, the circulating fluidized bed reactor may alsoinclude a catalyst separation loop 414, as is also shown in FIG. 9. Inthis loop, a sidestream of slurry is removed from port 415 in the slurryregion of the reaction zone and passed through a catalyst filter 417 forremoval of catalyst 418. Fresh catalyst 419 may be added to the filteredreacting solution that is conveniently recycled to the reactor by mixingit with fresh feed stream 401 and recirculation stream 420 forintroduction into draft tube 403. The catalyst separation loop 414 maybe operated continuously or intermittently as needed (e.g., to purgecatalyst having diminished activity and/or selectivity) and obviates theneed to periodically shut down the reactor for replacement ofparticulate catalyst. However, the capacity of catalyst filter 417 neednot be nearly as great as the filters used for separation of catalystfrom the reaction slurry exiting a continuous stirred tank reactor asdescribed above. Thus, significant savings in capital, operating, andmaintenance expense can be realized.

Various modifications may be made to the fluidized bed reactor 400 shownin FIG. 9. For example, rather than sparging the O₂-containing gas intothe reaction mixture 406 at the bottom of the reactor 400, an ejectornozzle similar to that shown in FIG. 7 may be provided at the top of thereactor through which both the aqueous feed stream 401 and theO₂-containing gas are combined and discharged into the reaction mixture406. Alternatively, circulation of the reacting mixture containing theparticulate catalyst may be provided by an impeller rotated within thedraft tube 405 in a manner to draw the reacting mixture downward throughthe draft tube and into the lower region of the oxidation reaction zone.Moreover, the catalyst separation loop 414 may optionally be integratedinto the heat transfer recirculation loop 420. The oxidation reactionconditions and operating parameters for a circulating fluidized bedreactor are similar to those described above for oxidation reactionzones provided by stirred tank reactors.

A still further alternative embodiment of the invention is illustratedin FIG. 10 wherein the oxidation of N-(phosphonomethyl)iminodiaceticacid substrate to N-(phosphonomethyl)glycine product is conducted in adistributed reactor system comprising a plurality of reactors 800A,800B, 800C, . . . 800 n in which reacting mixture progresses in seriesfrom each reactor to the succeeding reactor in the series. Primary feedmixture 814 comprising N-(phosphonomethyl)iminodiacetic acid substrateis introduced into reactor 800A and supplemental feed mixture is dividedinto parallel component feed streams 802B, 802C, 802D . . . 802 n thatare distributed among the series of reactors. Each reactor receives acomponent of distributed supply of oxygen or other oxidizing agent viafeed lines 804A, 804B, 804C, . . . 804 n. Optionally, a heat exchanger806A, 806B, 806C, etc. is interposed between each succeeding reactor andthe immediately preceding reactor to remove heat of reaction from theintermediate reaction mixture 810A, 810B, 810C, etc. exiting theimmediately preceding reactor 800A, 800B, 800C, etc. and allow eachreactor to operate adiabatically if desired. Alternatively, a cooledrecirculation stream 808A, 808B, 808C, 808D, etc. can be returned toeach reactor to remove exothermic heat of reaction and provide coolingof the reacting mixture in the reactor. In each of the series ofreactors subsequent to the first reactor 800A, a combined inlet stream812B, 812C, 812D, . . . 812 n is the combination of the component feedstream 802B, 802C, 802D, etc., the intermediate reaction mixture exitingthe immediately preceding reactor 810A, 8101B, 810C, etc., minus anyrecirculation 808A, 808B, 808C, 808D, etc. and any recirculation stream808B, 808C, 808D, etc. Each of reactors 800A, 800B, etc. can assume anyof the configurations described herein, but is preferably in the form ofa reactor in which catalyst is retained (e.g., a fixed bed or fluidizedbed reactor). A final reaction product 810 n is withdrawn from the lastreaction zone 800 n of the reaction zone series.

Each component feed stream 802B, 802C, etc. of a distributed reactorsystem may be highly concentrated, thereby contributing to highproductivity of the process. In fact, component feed streams comprisinga dense N-(phosphonomethyl)iminodiacetic acid substrate slurry or pastecomponent feed streams can be used. In each succeeding reactor followingthe first reactor 800A, (e.g., reactor 800B) a slurry or paste componentfeed mixture can be introduced, though it is preferred, especially inthe case of fixed bed reactors, that the combination of component 802Bfeed composition, component 802B feed rate, composition and flow rate ofintermediate reaction mixture 810A exiting the immediately precedingreactor 800A (minus any recirculation 808A), any recirculation 808B ofintermediate reaction mixture from reactor 800B are such that thecombined inlet stream 812B is substantially free of substrate solids orN-(phosphonomethyl)glycine product solids. However, it will beunderstood by those skilled in the art that, in certain embodiments ofthe invention, the component feed and intermediate reaction mixtures canbe in slurry form throughout, for example, where the oxidation catalystis a homogeneous catalyst, or where a monolithic catalyst such as fixedbed in honeycomb form is utilized.

Although both an additional component feed stream 802B, 802C, etc. andoxidant are preferably introduced into each of the series of reactionzones 800B, 800C, etc. following the first reaction zone, it will beunderstood that in a particular application, it may be necessary ordesirable to supply a component reaction mixture only to some but notall of the successive reaction zones. In some instances, it may not benecessary to supply oxidant to all reaction zones, though in most cases,supply of oxidant to each zone is preferred.

The fixed bed and distributed feed embodiments of the invention areuniquely suited for conversion of water-soluble salts ofN-(phosphonomethyl)iminodiacetic acid to water-soluble salts ofN-(phosphonomethyl)glycine. Because of the generally high solubility ofalkali metal and amine salts, e.g., potassium, ammonium, isopropylamine,and alkanolamine salts, of both the substrate acid and the product acid,either a fixed bed or stirred tank reactor may be operated at muchhigher substrate and product concentrations than is feasible in the acidprocess wherein productivity is limited by relatively low solubility. Infact, in the case of the salts, a fixed bed process may be particularlyadvantageous because it can be operated without need for any filtrationor centrifugation operation, either for removal of crystalline productor for removal of catalyst. An N-(phosphonomethyl)glycine salt solutioncan be formulated with various excipients commonly used in thecommercial application of N-(phosphonomethyl)glycine and soluble withminimal further processing. To produce the desirable concentrates ofcommerce, only modest concentrating steps are required. Extensiveimpurity separation may not be required.

A stirred tank reaction system, especially a continuous stirred tankreaction system may be advantageous for synthesis of salts because ofthe more substantial reaction heat load associated with the oxidation ofhigh concentration of substrate, and the accompanying exothermicoxidation of relatively large proportions of C₁ by products such asformaldehyde and formic acid. A continuous stirred tank reactor offers asignificant advantage over a batch reactor in the utilization ofreaction heat to preheat aqueous feed to the reactor. Combinations of aprimary continuous stirred tank for initial conversion with a fixed bedfinishing reactor may also be advantageous.

A fixed bed substantially plug flow reactor nonetheless offersparticular advantages, especially where the catalyst bed comprises noblemetal on carbon, because the plug flow operation serves to promoteoxidation of C₁ by-products, a reaction which is essentially first orderin C₁ substrate. However, for the same reason, plug flow exacerbates theheavy heat load associated with oxidation of an aqueous feed mixturecontaining a high concentration of substrate salt. Although therecirculation reaction system of FIG. 8 may be used to establishadequate heat transfer, it has an unfavorable effect on the kinetics ofdestruction of formaldehyde and formic acid, though depending on therecirculation rate the effect on C₁ destruction may remain marginallysuperior to a fully back-mixed reactor.

Accordingly, in some instances it may be advantageous to conduct theoxidation reaction in a reaction system wherein the fixed bed is cooledby indirect transfer of heat to a cooling fluid comprising a heattransfer or process fluid flowing through a conduit within or in contactwith the catalyst bed. For example, the fixed bed may be disposed withinthe shell or tube side of a shell and tube heat exchanger, with thecooling fluid being passed through the other side of the exchanger. Inone such embodiment, the fixed bed may comprise multiple component bedsseparately disposed in the tubes of the heat exchanger, with the aqueousfeed mixture and oxidant being distributed among the component beds andthe cooling fluid flowing through the shell side of said heat exchanger.In an alternative embodiment, the fixed bed may be contained within theshell of the heat exchanger, baffles on the shell side optionally beingused to assure substantially plug flow of the liquid phase through thebed.

Alternatively, salts of N-(phosphonomethyl)glycine may be prepared in aseries of reactors separated by heat exchangers for cooling theintermediate reaction solution as described above. The distributed feedreaction system of FIG. 10 may be especially advantageous in dealingwith the heat load generated in the oxidation ofN-(phosphonomethyl)iminodiacetic acid salts toN-(phosphonomethyl)glycine salts. As noted, especially highproductivities may be achieved where the substrate and product are bothwater-soluble salts. For example, where the aqueous feed mixture maycontains at least about 15% by weight of the substrate salt, the finalreaction mixture may contain at least about 12% by weight of awater-soluble product salt; where the aqueous feed mixture contains atleast about 25% by weight of a water-soluble substrate salt, the finaloxidation reaction mixture may contain at least about 20% by weight of awater-soluble product salt; and where the aqueous feed mixture containsat least about 35% by weight of a water-soluble substrate salt, thefinal oxidation reaction mixture may contain at least about 28% byweight of a water-soluble product salt; all on an acid equivalent basis.In fact, even high product salt concentrations can be realized, inexcess of 35%, preferably in excess of 40% or even 50% by weight. Asdescribed above, the final reaction product may be the primary reactionmixture obtained in a single reactor, the primary product fraction of asingle recirculating fixed bed system as depicted in FIG. 8 or theeffluent of the last of a series of reactors as is further describedabove.

The final reaction product is preferably further concentrated by removalof water therefrom. For example, the final reaction mixture may beintroduced into a flash evaporation zone wherein the pressure is lowerthan the vapor pressure of the final oxidation mixture at thetemperature at which it exits the reactor, or the last of a series ofreactors. With relatively low expenditure of energy, sufficient watermay be removed from the final oxidation reaction product to produce aconcentrated solution containing at least about 40% by weight of awater-soluble salt of N-(phosphonomethyl)glycine on an acid equivalent

Typically, the concentration of N-(phosphonomethyl)glycine product inthe oxidation reaction mixture effluent exiting the reactor systems ofthe present invention may be as great as 40% by weight, or greater.Preferably, the N-(phosphonomethyl)glycine product concentration is fromabout 5 to about 40%, more preferably from about 8 to about 30%, andstill more preferably from about 9 to about 15%. The concentration offormaldehyde in the product mixture is preferably less than about 5000ppm, more preferably less than about 4000 ppm, still more preferablyless than about 2800 ppm by weight, and still even more preferably lessthan about 1500 ppm. The concentration of formic acid in the productmixture is preferably less than about 12,000 ppm, more preferably lessthan about 4000 ppm, still more preferably less than about 2000 ppm byweight, and still even more preferably less than about 1500 ppm. Theconcentrations of aminomethylphosphonic acid (AMPA),N-methyl-aminomethylphosphonic acid (MAMPA),N-methyl-N-(phosphonomethyl)glycine (NMG) in the product mixture arereadily controlled at each less than 9000 ppm, can usually be controlledat less than 4500 ppm, and often be maintained below 1500 ppm. It willbe understood that these concentrations of by-products are based on asingle pass operation in which the only feed to the reactor system is anaqueous mixture containing N-(phosphonomethyl)iminodiacetic acid or saltthereof as obtained from the phosphonomethylation of iminodiacetic acid.Where any recycle stream such as the decantate from an adiabaticcrystallizer as described below is introduced into the reactor system,the attendant recycle of by-products tends to increase the by-productcontent of the reaction product mixture.

Purifying and/or Concentrating the N-(phosphonomethyl)glycine Product

Another aspect of this invention relates to purifying and/orconcentrating the N-(phosphonomethyl)glycine product obtained in theoxidation reaction mixture effluent. The various improvements inN-(phosphonomethyl)glycine product recovery provided by the presentinvention have wide application and, for example, may be used to recoverN-(phosphonomethyl)glycine product from the oxidation reaction mixtureproduced by the various continuous oxidation reactor systems describedherein. However, this further aspect of the present invention is notlimited to such application or to use in conjunction with continuousoxidation reactor systems generally. As will be apparent to thoseskilled in the art, the strategies set forth herein may beadvantageously applied in recovering N-(phosphonomethyl)glycine productfrom oxidation reaction mixture effluents produced by other reactorsystems as well, including batch reactor systems.

The reaction mixture normally contains water and various impuritiesbesides the desired N-(phosphonomethyl)glycine product. These impuritiesmay include, for example, various by-products and unreacted startingmaterials such as unreacted N-(phosphonomethyl)iminodiacetic acidsubstrate, N-formyl-N-(phosphonomethyl)glycine, phosphoric acid,phosphorous acid, hexamethylenetetraamine, aminomethylphosphonic acid,N-methyl-aminomethylphosphonic acid, iminodiacetic acid, formaldehyde,formic acid, and the like. The value of the N-(phosphonomethyl)glycineproduct normally dictates maximal recovery of the product from thereaction mixture and also often provides incentive for recycling atleast a portion of the depleted reaction mixture to the oxidationreaction zone(s) for further conversion of unreacted substrate andrecovery of unrecovered product.

Commercial considerations also sometimes dictate that the concentrationof the N-(phosphonomethyl)glycine product in the commercially soldmixtures be significantly greater than the concentrations in thereaction mixtures that are typically formed using the above describedoxidation reaction systems, particularly where theN-(phosphonomethyl)glycine product is being used for agriculturalpurposes. For example, when using a heterogenous catalyst to make theN-(phosphonomethyl)glycine free acid at the more preferred operatingtemperatures (i.e., from about 95 to about 105° C.), the maximumconcentration of the N-(phosphonomethyl)glycine product in the reactionmixture is preferably no greater than about 9% by weight so that it willremain solubilized. Sometimes, however, it is desirable for thecommercially sold mixtures to have an N-(phosphonomethyl)glycineconcentration which is significantly greater.

Thus, after the N-(phosphonomethyl)glycine product has been formed and,if necessary, separated from the catalyst, it is typically preferable toconcentrate the product and separate the product from the variousimpurities in the oxidation reaction mixture.

Smith (in U.S. Pat. No. 5,087,740) describes one process for purifyingand concentrating an N-(phosphonomethyl)glycine product. Smith disclosespassing a reaction mixture containing N-(phosphonomethyl)glycine througha first ion exchange resin column to remove impurities that are moreacidic than the N-(phosphonomethyl)glycine, passing the effluent fromthe first ion exchange resin column through a second ion exchange resincolumn which adsorbs the N-(phosphonomethyl)glycine, and recovering theN-(phosphonomethyl)glycine by passing a base or strong mineral acidthrough the second ion exchange resin column.

Many other techniques for purifying and concentrating anN-(phosphonomethyl)glycine product include a crystallization step,wherein the N-(phosphonomethyl)glycine product is crystallized toseparate it from at least a portion of the remaining reaction mixture.

The product recovery processes illustrated in FIGS. 11-14A and describedbelow have particular application in the concentration and recovery ofproduct from oxidation reaction mixtures containingN-(phosphonomethyl)glycine product susceptible to crystallization, andespecially those containing N-(phosphonomethyl)glycine free acid. Theconcentrated N-(phosphonomethyl)glycine free acid is typically used inthe preparation of the other N-(phosphonomethyl)glycine products such asthose described above.

In a particularly preferred embodiment, at least a portion of the finalreaction mixture (preferably absent any catalyst, and particularlyabsent any heterogeneous catalyst or homogeneous catalyst thatco-crystallizes with the N-(phosphonomethyl)glycine product) isintroduced into a non-adiabatic heat-driven evaporative crystallizer,where heat is added to the oxidation reaction mixture to evaporate offwater from the reaction mixture and thereby concentrate and crystallizethe N-(phosphonomethyl)glycine product. The heat used in thenon-adiabatic crystallizer is normally derived from steam. Preferably,at least about 30%, more preferably at least about 50%, even morepreferably from about 80% to about 100%, still even more preferably fromabout 90% to nearly 100% of the water in the reaction mixture isevaporated in the non-adiabatic crystallizer system. Evaporativecrystallization is particularly advantageous because it also separatesthe product from small molecule impurities, most notably formaldehydeand formic acid, which tend to evaporate from the reaction mixture alongwith the water.

The pressure in the heat-driven evaporative crystallizer preferably isno greater than about 10 psia, more preferably from about 1 to about 10psia, even more preferably from about 1 to about 5 psia, still morepreferably from about 2 to about 3 psia, and still yet even morepreferably about 2.8 psia. The operating temperature of the heat-drivenevaporative crystallizer preferably is no greater than about 80° C.,more preferably from about 40° C. to about 80° C., even more preferablyfrom about 50° C. to about 70° C., and still even more preferably about60° C.

FIG. 11 shows an example of one system which employs an evaporativecrystallizer. An aqueous feed 201 comprising anN-(phosphonomethyl)iminodiacetic acid substrate is introduced into anoxidation reactor system 203 comprising one or more oxidation reactionzone(s) wherein the substrate is oxidized to form an oxidation reactionmixture 205 comprising N-(phosphonomethyl)glycine product. In FIGS. 11,12, 13 and 14A, details of the oxidation reactor system, includingcatalyst separation and recycle mechanisms (e.g., catalyst filters,catalyst holding tanks, pre-filter flash tanks and the like) that may bepresent have been omitted, it being understood that the oxidationreaction mixture withdrawn from the reactor system has beensubstantially freed of catalyst, as necessary, depending on the specificreactor configuration(s) employed. The oxidation reaction mixture 205may optionally be passed through a pre-crystallizer flash tank 206. Thepre-crystallizer flash tank 206 lowers the pressure on the reactionmixture 205 to some degree causing dissolved CO₂ to be flashed out ofthe mixture and vented from the flash tank. An oxygen source (e.g., anO₂-containing gas) may be introduced into the pre-crystallizer flashtank 206 to further oxidize N-(phosphonomethyl)iminodiacetic acidsubstrate in the reaction mixture 205 that did not oxidize in theoxidation reaction zone(s) of the reactor system 203, as well as tofurther oxidize formaldehyde and formic acid by-products present in thereaction mixture 205. In this manner, the pre-crystallizer flash tank206 acts as an oxidation reaction zone in series with the reactor system203.

An evaporative crystallizer feed stream 239 is then introduced into theheat-driven evaporative crystallizer 207 in which heat is transferred tothe evaporative crystallizer feed stream 239 to vaporize water (andsmall molecule impurities, such as formaldehyde and formic acid) to formthe non-adiabatic crystallizer overhead vapor stream 209. A largeportion of the N-(phosphonomethyl)glycine product precipitates(typically from about 50% to about 60% on a per pass basis) to producean evaporative crystallization slurry 211. Slurry 211 is withdrawn fromthe non-adiabatic evaporative crystallizer 207, and can be introducedinto a hydroclone (or bank of hydroclones) 213, which forms aconcentrated slurry 215 enriched in precipitatedN-(phosphonomethyl)glycine product and a solids-depleted stream 221. Theconcentrated slurry 215 is introduced into a solids separation device,preferably a centrifuge, which forms a centrate 223 (which is furtherdepleted in precipitated N-(phosphonomethyl)glycine product) and anN-(phosphonomethyl)glycine product wet cake 219.

Normally, the concentration of the N-(phosphonomethyl)glycine product inthe wet cake 219 is at least about 95% (by weight of all compoundsbesides water). A lower product concentration may be tolerated if thewet cake 219 is subsequently washed with water or blended with higherpurity product as described below.

At least a portion of the heat-driven crystallizer overhead 209 may berecycled back to the oxidation reaction zone(s) of the reactor system203. In the embodiment shown in FIG. 11, a portion 243 is condensed andrecycled back for use as a source of water for dissolving theN-(phosphonomethyl)iminodiacetic acid substrate to form the feed stream201 for the reactor system 203. Preferably, condensate 243 is introducedinto the most upstream oxidation reaction zone where the reactor system203 comprises two or more oxidation reaction zones in series. Stream243, as with nearly all recycle streams of this invention, mayalternatively (or additionally) be introduced directly into theoxidation reaction zone(s) rather than combined with other ingredients(e.g., in aqueous feed stream 201) before being introduced into theoxidation reaction zone(s). Particularly where the catalyst is acarbon-containing catalyst and more particularly where the catalystcomprises carbon-supported noble metal, a portion of the non-adiabaticcrystallizer overhead 209, may also advantageously be used to reduce thecatalyst surface. This is due to the fact that the heat-drivenevaporative crystallizer overhead 209 typically contains formaldehydeand/or formic acid, which both act as reducing agents, particularlytoward carbon-containing catalysts and more particularly towardcatalysts comprising carbon-supported noble metal. Typically, theportion of the non-adiabatic crystallizer overhead 209 used in such areduction treatment is first condensed and the condensate may beintroduced into one or more catalyst holding tank(s) within reactorsystem 203 where the reduction treatment takes place. In addition toreducing the catalyst, such treatment may act to wash the catalyst andtakes advantage of the residence time of the catalyst in the catalystholding tank(s). In one particularly preferred embodiment, a portion ofthe non-adiabatic crystallizer overhead 209 is further rectified ordistilled to obtain a vapor stream containing an enriched concentrationof formaldehyde and/or formic acid. This enriched vapor stream, in turn,may be condensed and contacted with the carbon-containing catalyst.

At least another portion 241 of the heat-driven evaporative crystallizeroverhead 209 is typically purged (i.e., discharged) from the system aspurge stream 241. In a continuous system, this purge 241 helps to reducethe amount of waste buildup (particularly small molecule impuritybuildup) in the system and helps manage the water balance of the system.The purged waste 241 may, in turn, be further treated to removeimpurities. Such a treatment may include, for example, contacting thepurge stream 241 with an O₂-containing gas and a catalyst comprising aGroup VIII metal (particularly platinum, palladium, and/or rhodium) and,optionally, a carbon support, thereby oxidizing formaldehyde and formicacid to form environmentally benign CO₂ and water. The reaction ispreferably conducted at a temperature of from about room temperature toabout 90° C. (more preferably from about 50° C. to about 90° C.), apressure of from about atmospheric to about 200 psi, a dissolved oxygenconcentration of from about 1 to about 7 ppm, and a Group VIII metal toworking reactor volume ratio of from about 0.00015:1 to about 0.00040:1.This process is described in detail by Smith in U.S. Pat. No. 5,606,107.The product resulting from oxidation of the heat-driven evaporativecrystallizer overhead 209 may be recycled to the oxidation reactionzone(s) of reactor system 203 and used as a source of makeup water.

The hydroclone solids-depleted stream 221 is preferably recycled back tothe heat-driven evaporative crystallizer 207 for further recovery of theN-(phosphonomethyl)glycine product.

At least a portion 231 of the centrate 223 from the centrifuge 217 ispreferably recycled back to the heat-driven crystallizer 207 for furtherrecovery of the N-(phosphonomethyl)glycine product. Alternatively (or inaddition), a portion 233 of the centrate 223 can be recycled back to theoxidation reaction zone(s) of the reactor system 203 to convertunreacted N-(phosphonomethyl)iminodiacetic acid substrate in thecentrate 223 to N-(phosphonomethyl)glycine product. Alternatively (or inaddition), a portion 227 of the centrate 223 can be purged from thesystem.

Purging a portion 227 of the centrate 223 from the centrifuge 217 in acontinuous system helps to reduce the amount of impurity buildup(particularly larger molecule impurity buildup) in the system and thusin wet cake 219. Such a treatment may include, for example:

-   -   1. The purge stream 227 may be contacted with O₂ and a Group        VIII metal catalyst to oxidize formaldehyde and formic acid in        the purge stream 227, as described above for the non-adiabatic        crystallizer overhead purge 241.    -   2. The purged waste 227 may be contacted with O₂ and a        noble-metal-containing catalyst to oxidatively cleave any        N-substituted-N-(phosphonomethyl)glycine (often most notably        N-methyl-N-(phosphonomethyl)glycine) to form additional        N-(phosphonomethyl)glycine product, which, in turn, may be        collected in a crystallizer, such as by recycling it back to the        non-adiabatic crystallizer 207. Preferably, this reaction is        conducted at a pressure of at least atmospheric pressure (more        preferably from about 30 to about 200 psig), a temperature of        from about 50° C. to about 200° C. (more preferably from about        70° C. to about 150° C., and even more preferably from about        125° C. to about 150° C.), a dissolved oxygen concentration of        no greater than about 2 ppm, and a weight ratio of the noble        metal to the N-substituted-N-(phosphonomethyl)glycine        by-product(s) of from about 1:500 to about 1:5 (more preferably        from about 1:200 to about 1:10, and even more preferably from        about 1:50 to about 1:10). This method of treatment is described        in detail by Morgenstern et al. in U.S. Pat. No. 6,005,140.    -   3. The purged waste 227 may be combined with formaldehyde in        stoichiometric excess relative to the N-(phosphonomethyl)glycine        compounds and derivatives thereof, and then heated in the        presence of a transition metal catalyst (e.g., manganese,        cobalt, iron, nickel, chromium, ruthenium, aluminum, aluminum,        molybdenum, vanadium, copper, zinc, or cerium) to form more        environmentally benign compounds. This process is described in        detail by Grabiak et al. in U.S. Pat. No. 4,851,131.    -   4. The purged waste 227 may be passed through another        crystallizer for further recovery of N-(phosphonomethyl)glycine        product.

In another particularly preferred embodiment, at least a portion of theoxidation reaction mixture effluent (preferably absent any catalyst,particularly absent any heterogeneous catalyst or a homogeneous catalystthat co-precipitates with the N-(phosphonomethyl)glycine product) isintroduced into a crystallizer which operates substantiallyadiabatically (i.e., any heat input or removal to the crystallizer is nogreater than about 200 kcal. per kg of oxidation reaction mixture fed tothe crystallizer), and more preferably fully adiabatically. Unlike theprocess as conducted in a non-adiabatic crystallizer as described above,the separation process in an adiabatic crystallizer results primarilyfrom reduction in solubility due to cooling rather than to theconcentrating effect of removal of water. In a preferred embodiment,separation of mother liquor from precipitated crystallization solids isaccomplished in part by decantation. Because the amount of water removedin adiabatic crystallization is relatively small, the mother liquor hasa relatively low impurities content. In accordance with the invention,it has been discovered that this mother liquor may be directly recycledto the oxidation reactor system as a source of process water. Adiabaticcrystallization is advantageous because it does not require the energy(typically in the form of steam) that is required for the evaporation ina non-adiabatic crystallizer.

In an especially preferred adiabatic crystallizer system, the finalreaction mixture is subjected to a sudden drop in pressure in a flashsection, which causes part of the water in the reaction mixture toevaporate. This evaporation, in turn, causes the remaining reactionmixture to cool. Cooling results in the precipitation ofN-(phosphonomethyl)glycine product. Mother liquor may then be decantedto concentrate the slurry of the N-(phosphonomethyl)glycine product.Adiabatic crystallization is advantageous because it does not requirethe energy (typically in the form of steam) that is required for theevaporation in a non-adiabatic crystallizer.

FIG. 12 shows one embodiment of a system comprising an adiabaticcrystallizer 115. An aqueous feed 101 comprising anN-(phosphonomethyl)iminodiacetic acid substrate is introduced into anoxidation reactor system 103 comprising one or more oxidation reactionzone(s) wherein the substrate is oxidized to form an oxidation reactionmixture 105 comprising N-(phosphonomethyl)glycine product. The oxidationreaction mixture 105 withdrawn from the reactor system 103 mayoptionally be passed through a pre-crystallizer flash tank 107. Thepre-crystallizer flash tank 107 lowers the pressure on the reactionmixture 105 to some degree causing dissolved CO₂ to be flashed out ofthe mixture and vented from the flash tank. The preferred pressure dropdepends on the pressure at which the oxidation reaction is conducted inthe reactor system 103. Normally, for example, where the oxidationreaction zone(s) pressure is 115 psia, the pressure drop in thepre-crystallizer flash tank 107 is no greater than about 100 psig, morepreferably from about 20 to about 80 psig, even more preferably fromabout 60 to about 80 psig, and still even more preferably about 75 psig;while the preferred pressure drop where the reaction zone(s) pressure is215 psia is no greater than about 200 psig, more preferably from about120 to about 180 psig, even more preferably from about 160 to about 180psig, and still even more preferably about 175 psig. This typicallycauses up to about 1.5% (more typically from about 0.2 to about 1%, evenmore typically from about 0.2 to about 0.5%, and still even moretypically about 0.25%) by weight of the final reaction mixture 105 to gointo the vapor phase. Typically, the pressure over the resultingcrystallizer feed steam 114 leaving the pre-crystallizer flash tank 107is at least about 15 psia, more preferably from about 25 to about 100psia, even more preferably from about 30 to about 60 psia, and stilleven more preferably about 40 psia.

An oxygen source (e.g., an O2-containing gas) may be introduced into thepre-crystallizer flash tank 107 to further oxidizeN-(phosphonomethyl)iminodiacetic acid substrate in the reaction mixture105 that did not oxidize in the oxidation reactor system 103, as well asto further oxidize formaldehyde and formic acid by-products present inthe reaction mixture 105. In this manner, the pre-crystallizer flashtank 107 acts as an oxidation reaction zone in series with the reactorsystem 103.

The crystallizer feed stream 114 is introduced into the adiabaticcrystallizer 115. A detailed description of the operation of anadiabatic crystallizer system in accordance with the present inventionis set forth below in connection with FIG. 12A. Operation of theadiabatic crystallizer 115 produces vapor 117 (i.e., the adiabaticcrystallizer overhead) discharged from the top of the crystallizer, adecantate (i.e., mother liquor) stream 124 withdrawn from thecrystallizer and a crystallization product slurry 125 comprisingprecipitated crystalline N-(phosphonomethyl)glycine product removed fromthe bottom of the crystallizer.

At least a portion 146 of the adiabatic crystallizer overhead 117 and/orand at least a portion 132 of the withdrawn decantate 124 may berecycled back to the oxidation reaction zone(s) of the reactor system103. Typically, the recycled adiabatic crystallizer overhead 117 and/orwithdrawn decantate 124 is/are recycled back to the oxidation reactionzone(s) and used as a source of water for dissolving theN-(phosphonomethyl)iminodiacetic acid substrate to form the feed stream101 for the reactor system 103. Preferably, the recycled adiabaticcrystallizer overhead 117 and/or withdrawn decantate 124 is/areintroduced into the most upstream oxidation reaction zone where thereactor system 103 comprises two or more oxidation reaction zones inseries. Recycling at least a portion 132 of the decantate 124 back tothe reactor system is advantageous because it reduces the waterrequirements and the volume of waste from the system. It also oftenallows recovery of additional N-(phosphonomethyl)glycine product fromthe unreacted N-(phosphonomethyl)iminodiacetic acid substrate in thedecantate 124. This recycle is additionally advantageous because itoften allows for additional by-products, such as formaldehyde and formicacid, to be oxidized. The recycle of stream 132 is further advantageousbecause it allows water to be recycled directly back to the oxidationreaction zone(s) from the crystallizer 115 without having to expendenergy to evaporate the water (as is the case with the recycle of theoverheads from the heat-driven evaporative crystallizer discussedabove). Because the recycled decantate 132 also remains at a relativelyelevated temperature (most preferably 60° C.), the recycled decantate132 can be used to preheat the aqueous feed stream 101. When a noblemetal on carbon catalyst is utilized in the reactor system 103, a stillfurther benefit of decantate recycle stream 132 may be realized in thatnoble metal leached from the catalyst is returned to the reactor system.It is believed that recycling noble metal-containing streams such asstream 132 to the reactor system 103 reduces the net loss of noble metalfrom the system. A portion of the leached noble metal contained in sucha recycle stream may redeposit on the surface of the heterogeneouscatalyst in the catalytic reactor system.

Particularly where the catalyst is a carbon-containing catalyst (andeven more particularly where the catalyst comprises a carbon-supportednoble metal), it is preferable to recycle at least a portion of theadiabatic crystallizer overhead 117 indirectly by condensing it and thenmixing the condensate with the catalyst. This is often advantageousbecause the adiabatic crystallizer overhead 117 often containsformaldehyde and/or formic acid, which, as noted above, both act asreducing agents. In one particularly preferred embodiment, a portion ofthe adiabatic crystallizer overhead 117 is rectified or condensed andfurther distilled to obtain a condensate enriched in formaldehyde and/orformic acid. This enriched solution, in turn, is the portion of theadiabatic crystallizer overhead that is contacted with thecarbon-containing catalyst. As noted above, this reduction treatment canoccur in one or more catalyst holding tank(s) within the reactor system103.

At least another portion 149 of the adiabatic crystallizer overhead 117and/or at least a portion 151 of the withdrawn decantate 124 may bepurged (i.e., discharged) from the system as waste. In a continuoussystem, this purge helps to reduce the amount of impurity buildup in thesystem. This purged waste may, in turn, be further treated to removeimpurities by techniques known in the art, such as those described abovefor the purged waste stream of the centrifuge downstream of anon-adiabatic crystallizer. For example, the purged waste may becontacted with an O₂-containing gas and Group VIII metal catalyst tooxidize formaldehyde and formic acid to CO₂ and water. The product ofsuch oxidation treatment may be recycled to the oxidation reactionzone(s) of the reactor system 103 and used as a source of makeup water.

The N-(phosphonomethyl)glycine product slurry 125 withdrawn from thebottom of the adiabatic crystallizer 115 contains the bulk of theN-(phosphonomethyl)glycine product. The slurry 125 is typically passedthrough a centrifuge 155 to further concentrate the slurry 125 and forma wet cake 157 containing the N-(phosphonomethyl)glycine product.Normally, the concentration of the N-(phosphonomethyl)glycine product inthe wet cake 157 is at least about 95% (by weight of all compoundsbesides water). The solids-depleted stream 161 (i.e., the centrate) fromthe centrifuge 155 may, for example, be recycled back to the adiabaticcrystallizer 115 via stream 165 or recycled back to the oxidationreaction zone(s) of the reactor system 103 via stream 169 to be used asa source of water in the aqueous feed stream 101. In order to maintainimpurity concentrations at acceptable levels and enable the advantageoususe of recycled decantate stream 132, at least a portion of thesolids-depleted stream 161 may be removed via stream 173. Stream 173 maybe subsequently treated by, for example, the waste treatment processesdescribed above for the purge stream of the centrifuge downstream of anon-adiabatic heat-driven crystallizer. In a further embodiment, stream173 is sent to a heat-driven evaporative crystallizer for additionalproduct recovery in manner similar to that shown in FIG. 13.

FIG. 12A is a schematic of a preferred adiabatic crystallization systemfor use in the practice of the present invention. As shown, the system115 comprises vapor/liquid separator 703 defining a vapor/liquidseparation zone positioned generally above and in fluid flowcommunication with a retention chamber 705. Vapor/liquid separator 703is segregated from direct communication with the upper region ofretention chamber 705, but is in fluid flow communication with the lowerregion of the retention chamber via a draft tube 706, the mouth 708 ofdraft tube 706 being separated by only a relatively short distance fromthe bottom of the retention chamber. Crystallizer apparatus of thisgeneral configuration are available from HPD Products Division of U.S.Filter (Plainfield, Ill., U.S.A.). A crystallizer recirculation inlet709 is located on vapor/liquid separator 703, while retention chamber705 is provided with a decantation liquid exit 711 for crystallizationmother liquor located above mouth 708 of draft tube 706, an intermediaterecirculation slurry exit 712 located above the mouth of the draft tubeand below decantation liquid exit 711 and a lower product slurry exit713 located at the bottom of chamber 705. During operation, retentionchamber 705 is essentially filled with liquid while the liquid level 715in vapor/liquid separator 703 is maintained somewhat below crystallizerrecirculation inlet 709.

An aqueous crystallizer feed mixture 716 comprisingN-(phosphonomethyl)glycine product obtained from the reaction mixtureeffluent 114 withdrawn from the oxidation reaction zone(s) of theoxidation reactor system (along with various recycle streams as will bedescribed below) is introduced through recirculation inlet 709 intovapor/liquid separator 703. The vapor/liquid separator defines anevaporation zone maintained by a vacuum system (not shown) atsub-atmospheric pressure and below the vapor pressure of thecrystallizer feed mixture 716. The liquid level 715 in the vapor/liquidseparator 703 is maintained by pressure equilibration through holesprovided in the upper section of the draft tube 708 communicating withthe retention chamber 705. The crystallizer feed mixture 716 comprises:(a) the oxidation reaction mixture effluent 114 withdrawn from theoxidation reaction zone(s) of the reactor system (i.e., startingsolution) which may have been filtered to remove catalyst; (b) a recycleslurry stream 723 comprising at least a portion of the slurry exitingthe intermediate recirculation slurry exit 712 as described below; andtypically also (c) a centrate 165 comprising crystallization motherliquor recycled from a centrifuge system 155 to which a crystallizationproduct slurry 125 from exit 713 is directed for recovery of solidN-(phosphonomethyl)glycine product, as further described below. Thepressure maintained in vapor/liquid separator 703 is generally nogreater than about 8 psia, preferably from about 1.5 to about 4 psia,even more preferably from about 2.5 to about 3.5 psia, and still evenmore preferably about 3 psia. Typically, the pressure of thecrystallizer feed mixture 716 immediately upstream of the vapor/liquidseparator is such that the feed mixture is subjected to a pressurereduction of at least about 100 psig, preferably from about 10 to about80 psig, more preferably from about 30 about 60 psig, and even morepreferably of about 38 psig, upon entry into the vapor/liquid separator.The sudden decrease in pressure causes water and small moleculeimpurities (e.g., formaldehyde and formic acid) to flash (i.e.,evaporate) from the feed mixture 716 in the vapor/liquid separator 703.The vapor 117 (i.e., overhead) produced is separated and discharged fromthe top of separator 703 and directed to a condensing unit (not shown).Normally, no greater than about 30% by weight, more preferably fromabout 5% to about 30% by weight, and even more preferably from about 5%to about 10% by weight of the oxidation reaction mixture 114 isdischarged as vapor 117. As a result of evaporation, the remainingcondensed phase portion of the crystallizer feed mixture 716 is cooledconsiderably, thereby resulting in precipitation ofN-(phosphonomethyl)glycine product and producing an evaporation productslurry 718 comprising crystalline N-(phosphonomethyl)glycine productsolids 719 suspended in mother liquor that is substantially saturated orsupersaturated in N-(phosphonomethyl)glycine product. Preferably, thecooling effect resulting from the pressure reduction enteringvapor/liquid separator 703 is sufficient that the temperature of theevaporation product slurry 718 is from about 30° C. to about 40° C.lower than the temperature of the oxidation reaction mixture 114introduced into the adiabatic crystallization system. The temperature ofthe evaporation product slurry 718 is no greater than about 80° C., morepreferably from about 45° C. to about 80° C., even more preferably fromabout 55° C. to about 70° C., and especially from about 60° C. to about70° C.

The evaporation product slurry 718 exits separator 703 by descendingdraft tube 706 and is introduced into the lower region of a retentionzone within retention chamber 705. The retention zone is divided into alower crystallization region, generally below level 720, and an upperdecantation region, generally above level 720. In the retention zone,the evaporation product slurry 718 is separated into a supernatantliquid 722 comprising a fraction of the mother liquor (typically the netproduction thereof) and second slurry stream 723 comprising precipitatedN-(phosphonomethyl)glycine product crystals and mother liquor which iswithdrawn from the retention chamber 705 through intermediate slurryexit 712. A decantate stream 124 comprising supernatant liquid 722 iswithdrawn from retention chamber 705 through decantation exit 711 nearthe top of the retention chamber in the decantation region.

Crystallization product slurry 125 comprising N-(phosphonomethyl)glycineproduct slurry is withdrawn from the bottom of retention chamber 705through exit 713 in the crystallization region. The crystallizationproduct slurry is forwarded to centrifuge system 155 whereinN-(phosphonomethyl)glycine product crystals are separated as wet cake.Normally N-(phosphonomethyl)glycine product in the wet cake is at leastabout 95% (by weight of all compounds besides water). The resultingcentrate 165 is recycled and combined with a second portion of productslurry in slurry steam 723 withdrawn from retention chamber 705 at theinterface (i.e., cloud zone) between the decantation region and thecrystallization region of the retention zone. The combined flow isintroduced into vapor/liquid separator 703 along with the oxidationreaction mixture 114 as the crystallizer feed mixture 716.

At least a major portion, preferably substantially all of the secondslurry stream 723 withdrawn from exit 712 is recirculated to thevapor/liquid separator 703, being mixed with the reaction mixture stream114 and the centrate 165 from centrifuge 155 to form the feed mixture716 to the vapor/liquid separator. Mother liquor 722 is separated (i.e.,decanted) from the evaporation product slurry 718 in the decantationregion. Decantation is accomplished by maintaining the relative rates atwhich reaction mixture 114 is introduced through inlet 709, decantate124 is withdrawn from exit 711, and all or a portion of the secondslurry 723 is recirculated from intermediate slurry exit 712 viacrystallizer feed stream 716 (thereby controlling the rate at whichevaporation product slurry 718 is introduced into the retention zone)such that the upward flow velocity in the lower crystallization regionof the retention zone below the intermediate slurry exit 712 issufficient to maintain precipitated N-(phosphonomethyl)glycine productcrystals 719 in suspension (i.e., entrained) in the liquid phase, whilethe upward flow velocity in the upper decantation region of theretention zone above the intermediate slurry exit 712, is below thesettling velocity of at least about 80% by weight, preferably below thesettling velocity of at least about 95% by weight, most preferably belowthe settling velocity of at least about 98% by weight, of theN-(phosphonomethyl)glycine product crystals 719 in the crystallizationregion. Thus, an interface is established, at approximately the level ofintermediate slurry exit 712, between the upper decantation region ofthe retention zone which contains substantially clear mother liquor, andthe lower crystallization region of the retention zone which contains acrystallization slurry.

Preferably, the relative rates at which oxidation reaction mixture 114is introduced into the adiabatic crystallization system 115, decantate124 is withdrawn from exit 711, product slurry 125 is withdrawn fromexit 713, and centrate 165 is recycled from centrifuge 155 arecontrolled so that the ratio of N-(phosphonomethyl)glycine productsolids to mother liquor in the lower crystallization region of theretention zone is higher than the incremental ratio ofN-(phosphonomethyl)glycine product to mother liquor ratio resulting fromthe effects of the evaporation, such incremental ratio being the ratioof N-(phosphonomethyl)glycine product solids incrementally produced tothe mother liquor incrementally produced thereby, i.e., the netproduction of crystalline N-(phosphonomethyl)glycine product. Expressedanother way, the incremental ratio is the ratio that would be obtainedif the oxidation reaction mixture were flashed in the absence of solidsin the crystallizer feed mixture (i.e., in the absence of recirculatedsecond slurry). It will be understood that the effects of evaporationinclude both the concentrating effects and cooling effects; but whereoperation of the crystallizer is substantially adiabatic, as ispreferred, crystallization results primarily from cooling of the liquidphase to a temperature at which solubility of N-(phosphonomethyl)glycineproduct is substantially lower than it is at the temperature of theoxidation reaction mixture. Preferably, the solids/mother liquor ratioin the lower region of the retention zone is at least about twice theincremental ratio resulting from evaporation effects, and theconcentration of product solids in the crystallization region is also atleast twice the concentration incrementally produced. Expressed inanother way, the N-(phosphonomethyl)glycine product solids concentrationin the lower crystallization region of the retention zone is at leastabout 12% by weight, preferably at least about 15% by weight, morepreferably in the range of between about 18% and about 25% by weight. Inthe system as illustrated, the rate of removal of solid product in thecentrifuge 155 and of mother liquor as decantate 124 are ultimatelyfixed by the system material balance, but the solids inventory in thelower crystallization region of the retention zone may be adjusted bytransient decrease or increase of the rate at which product slurry isremoved from exit 713 relative to the decantation rate via exit 711. Asis further discussed hereinbelow, control of the solids inventory in thecrystallization region of the retention zone has been found to providecontrol of the average particle size of the N-(phosphonomethyl)glycineproduct of the crystallization process.

The steady state upward flow velocity in the upper decantation region ofthe retention zone is determined by sizing the cross section ofretention chamber 705 based on the process material balance and thesolids settling velocity. Preferably, a relatively high upward velocityin the lower crystallization region of the retention zone is establishedby maintaining a high rate of recirculation of the second slurry 723between intermediate slurry exit 712 and the recirculation inlet 709 tothe vapor/liquid separator 703 (e.g., in the range of 20:1 to 100:1relative to the rate of oxidation reaction medium 114 introduced intothe crystallization system or decantate 124 removed therefrom). Thefraction of centrate 165 from centrifuge 155 that is recycled as part ofcrystallizer feed mixture 716 also augments the rate of recirculationand upward flow velocity, but otherwise tends to dilute the slurry inthe crystallization region. By combining a high second slurryrecirculation rate with proper sizing of retention chamber 705, theupward flow velocity in the lower crystallization region of theretention zone can be controlled at least four times the sedimentationvelocity of at least 80% by weight of the solids contained therein whilethe upward flow velocity in the upper decantation region of theretention zone can be established at less than one fourth of thesedimentation velocity of at least 80% by weight of the solids containedin the second slurry stream.

Operating at a high solids content in the crystallization region,combined with a high rate of recirculation of second slurry 723 fromintermediate slurry exit 712 to vapor/liquid separator 703 furtherprovides a high solids concentration throughout the evaporation zone.This mode of operation has been found to have a significantly favorableeffect on both productivity of the crystallization process and theparticle size and drainage characteristics of the crystallineN-(phosphonomethyl)glycine product obtained. There is a trade offbetween particle size and productivity because productivity relatespositively to the degree of supersaturation, but the degree ofsupersaturation generally correlates negatively with particle size.However, even at the relatively high average particle size, operation athigh solids content effectively increases the surface of solidN-(phosphonomethyl)glycine product crystals on which crystallization canoccur, and thus allows the crystallization process to proceed at highproductivity with relatively minimal supersaturation of the liquid phaserequired to provide a driving force for crystallization. Crystallizationat low levels of supersaturation in turn promotes formation ofrelatively large crystals. Thus, for a given productivity of theevaporation zone, the crystallization process of the invention providesan average particle size substantially larger than the average particlesize obtained in a reference evaporator that is fully back mixed andwherein the ratio of N-(phosphonomethyl)glycine product solids to motherliquor is equal to the ratio of N-(phosphonomethyl)glycine productsolids incrementally produced by the effects of the evaporation to themother liquor incrementally produced thereby. For example, thecrystallization process of the invention can be operated at highproductivity to obtain a product having (1) a median cube weightedparticle size of at least about 200 μm, preferably at least about 225μm, more preferably at least about 250 μm, even more preferably at leastabout 275 μm, still more preferably at least about 300 μm, andespecially at least about 350 μm; (2) a median length weighted particlesize of at least about 85 μm, preferably at least about 90 μm, morepreferably at least about 95 μm, even more preferably at least about 100μm, still more preferably at least about 105 μm, and especially at leastabout 110 μm; and (3) a BET (Brunauer-Emmett-Teller) surface area notgreater than about 0.11 m²/g, more preferably not greater than about0.09 m²/g, even more preferably not greater than about 0.07 m²/g, andstill even more preferably not greater than about 0.06 m²/g. The mediancube weighted and median length weighted particle sizes set forth abovemay be determined using a focused beam reflectance measurement (FBRM)device such as a LASENTEC Model M100F available from Laser SensorTechnology Inc. (Redmond, Wash., U.S.A.).

At the high flow rates prevailing along the recirculation path betweenintermediate slurry exit 712 and mouth 708 of draft tube 706, thecrystallization system operates in an essentially plug flow manner(i.e., without substantial axial back-mixing). As a result, a descendinggradient in the degree of supersaturation prevails along this path,thereby maximizing the integrated average driving force available forcrystallization, and enabling a lower degree of supersaturation to berealized at the downstream end of the plug flow path (i.e., at the mouthof the draft tube) than could be realized in a back-mixed system.Compounding the effect of plug flow with the generally low degree ofsupersaturation made feasible by the high crystal surface area presentedby the high solids content within the recirculating slurry, the netresult at any given rate of production is to further reduce the degreeof supersaturation in the liquid phase within the lower crystallizationregion of the retention zone, and therefore in the decantate motherliquor 124 that is removed from the system.

Plug flow operation in combination with high solids content can also beexploited with respect to productivity. It has been found that highproductivity can be realized even where the maximum supersaturation,i.e., the driving force for crystallization, expressed as the differencebetween the N-(phosphonomethyl)glycine product concentration in theaqueous liquid phase at any location within the recirculation path andthe saturation concentration of N-(phosphonomethyl)glycine product inthe aqueous liquid phase at such location is not greater than about 0.7%by weight, basis the aqueous liquid phase; or where the integratedaverage extent of supersaturation over the recirculation path is notgreater than 0.5%. Looking at the relationship between supersaturationand productivity in yet another way, the process as described canoperate effectively at an integrated average supersaturation over therecirculation path that is at least 0.2% lower than the extent ofsupersaturation required to provide equivalent crystallizationproductivity per unit working volume of a reference evaporatorconsisting of a fully back mixed evaporation zone in which the ratio ofN-(phosphonomethyl)glycine product solids to mother liquor is equal tothe ratio of N-(phosphonomethyl)glycine product solids incrementallyproduced by the effects of the evaporation to mother liquorincrementally produced thereby.

Because of the coarse particle size of the crystals produced inaccordance with the process illustrated in FIG. 12A, the capacity of acentrifuge or filter for dewatering product slurry 125 is substantiallyenhanced, with attendant savings in capital and maintenance expense. Forexample, by use of a vertical basket centrifuge system or otherdewatering device, a crystalline N-(phosphonomethyl)glycine product isobtained having a relatively low water content, e.g., exhibiting a losson drying of not greater than about 15% by weight, more preferably notgreater than about 8% by weight. Lower centrifuge cake wetnesstranslates directly into lesser contamination of the centrifuge cakewith chlorides, NMG, unreacted N-(phosphonomethyl)iminodiacetic acidsubstrate, etc. Thus, the use of the adiabatic crystallization systemaffords the opportunity to produce an exceptionally pure grade ofN-(phosphonomethyl)glycine product. So as to minimize attrition of theN-(phosphonomethyl)glycine product crystals, an axial flow pump is usedto recirculate material in the adiabatic crystallizer system.

Advantageously, the crystallization operation of FIG. 12A is conductedadiabatically, i.e., there is no substantial transfer of heat to or fromthe system via heat transfer to or from the vapor/liquid separationzone, the retention zone, the feed mixture to the vapor/liquidseparation zone, the second slurry that is recirculated to thevapor/liquid separation zone, or the centrate that is returned from thecentrifuge system. By reduction in pressure of the feed mixture asdescribed above, sufficient evaporation is achieved for cooling theliquid phase to the extent that substantial crystallization ofN-(phosphonomethyl)glycine product is realized. Capital and energysavings are realized by obviating the need for evaporator heatexchangers, and process downtime required to periodically parboil fouledheat exchangers is also eliminated.

Moreover, without substantial expenditure of energy for the separation,a decantation stream 124 is provided which can readily serve as a sourceof process water for recycle to the oxidation reaction zone(s) of theoxidation reactor system. Because crystallization can be effected athigh productivity at a relatively limited degree of supersaturation, thedecantate recycled to the oxidation reaction zone(s) has nearly theminimum theoretical N-(phosphonomethyl)glycine product concentration.Since productivity of the oxidation reaction system may typically belimited by the solubility of N-(phosphonomethyl)glycine product in thereaction mixture effluent, especially where a particulate catalyst isutilized in the preparation of the free acid form of N-(phosphonomethyl)glycine, N-(phosphonomethyl)glycine product in the recycle decantate canat least marginally detract from the net productivity of the oxidationreactor system by limiting the rate at whichN-(phosphonomethyl)iminodiacetic acid substrate can be converted toN-(phosphonomethyl)glycine product therein without exceedingN-(phosphonomethyl)glycine product solubility. This modest penaltyassociated with decantate recycle is minimized by recovering nearly themaximum theoretical proportion of N-(phosphonomethyl)glycine product inthe crystallizer and thereby minimizing the N-(phosphonomethyl)glycineproduct content of the water stream that is recycled.

Although the system depicted in FIG. 12A is preferred, those skilled inthe art will recognize that other options exist for establishing andmaintaining the high N-(phosphonomethyl)glycine product to mother liquorratios in the crystallization region of the retention zone that areeffective to provide relatively coarse crystals. The process of FIG. 12Ais effective to retain solids in the evaporation zone; the process couldalternatively be operated to return solids to the evaporation zone. Forexample, if crystallization is conducted in a fully back mixedevaporative crystallizer, it is possible to establish and maintain ahigh and disproportionate inventory of solids in the crystallizer byrecycling crystalline product from the filter or centrifuge used forN-(phosphonomethyl)glycine product solids recovery, while either notrecycling filtrate/centrate or recycling a lesser proportion offiltrate/centrate relative to solids recycled. However, as those skilledin the art will appreciate, the latter scheme of operation comes with apenalty in capital intensive filter or centrifuge capacity. Asignificant advantage of the preferred process of FIG. 12A is theachievement of high solids inventory by decantation rather than byexpensive mechanical means for solid/liquid separation.

Surprisingly, it has been discovered that operation of the adiabaticcrystallizer in the preferred, manner as described above, may obviatethe need for concentrating the product slurry of the decantation (as byuse of a hydroclone) prior to introduction of the slurry into acentrifuge.

FIG. 13 shows an example of one preferred embodiment wherein the processcomprises an adiabatic crystallizer operating in series with anon-adiabatic crystallizer.

Many of the various streams shown in FIG. 13 are analogous to thosedescribed above for non-adiabatic crystallizers and adiabaticcrystallizers alone. An aqueous feed stream 601 comprising anN-(phosphonomethyl)iminodiacetic acid substrate is introduced into anoxidation reactor system 603 comprising one or more oxidation reactionzone(s) wherein the substrate is oxidized to form an oxidation reactionmixture 605 comprising N-(phosphonomethyl)glycine product. The oxidationreaction mixture 605 may optionally be passed through a pre-crystallizerflash tank 607. The pre-crystallizer flash tank 607 lowers the pressureon the reaction mixture 605 to some degree causing dissolved CO₂ to beflashed out of the mixture and vented from the flash tank. An oxygensource (e.g., an O₂-containing gas) may be introduced into thepre-crystallizer flash tank 607 to further oxidizeN-(phosphonomethyl)iminodiacetic acid substrate in the reaction mixture605 that did not oxidize in the oxidation reaction zone(s) of thereactor system 603, as well as to further oxidize formaldehyde andformic acid by-products present in the reaction mixture 605. In thismanner, the pre-crystallizer flash tank 607 acts as an oxidationreaction zone in series with the reactor system 603.

A crystallizer feed stream 614 which comprises most of theN-(phosphonomethyl)glycine product is introduced into the adiabaticcrystallizer 615. Operation of the adiabatic crystallizer 615 producesvapor 617 (i.e., the adiabatic crystallizer overhead) discharged fromthe top of the crystallizer, a decantate (i.e., primary mother liquor)stream 624 withdrawn from the crystallizer and a primary crystallizationproduct slurry 625 comprising precipitated crystallineN-(phosphonomethyl)glycine product removed from the bottom of thecrystallizer.

At least a portion 646 of the adiabatic crystallizer overhead 617 and/orand at least a portion 632 of the withdrawn decantate 624 may berecycled back to the oxidation reaction zone(s) of reactor system 603.Typically, the recycled adiabatic crystallizer overhead 617 and/orwithdrawn decantate 624 is/are recycled back to the oxidation reactionzone(s) and used as a source of water for dissolving theN-(phosphonomethyl)iminodiacetic acid substrate to form the feed stream601 for the reactor system 603. Preferably, the recycled adiabaticcrystallizer overhead 617 and/or withdrawn decantate 624 is/areintroduced into the most upstream oxidation reaction zone where thereactor system 603 comprises two or more oxidation reaction zones inseries. Recycling at least a portion 632 of the decantate 624 back tothe oxidation reactor system is advantageous because it reduces thewater requirements and the volume of waste from the system. It alsooften allows recovery of additional N-(phosphonomethyl)glycine productfrom the unreacted N-(phosphonomethyl)iminodiacetic acid substrate inthe decantate 624. This recycle is additionally advantageous because itoften allows for additional by-products, such as formaldehyde and formicacid, to be oxidized. The recycle of stream 632 is further advantageousbecause it allows water to be recycled directly back to the oxidationreaction zone(s) from the crystallizer 615 without having to firstexpend energy to evaporate the water (as is the case with thenon-adiabatic heat-driven crystallizer discussed above). Because therecycled decantate 632 also remains at a relatively elevated temperature(most preferably 60° C.), the recycled decantate 632 can be used topreheat the aqueous feed stream 601.

Particularly where the catalyst is a carbon-containing catalyst (andespecially such a catalyst which also comprises a noble metal), it ispreferable to recycle at least a portion of the adiabatic crystallizeroverhead 617 indirectly by mixing it with the catalyst. This isadvantageous because the adiabatic crystallizer overhead 617 oftencontains formaldehyde and/or formic acid, which, as noted above, bothact as reducing agents. In one particularly preferred embodiment, aportion of the adiabatic crystallizer overhead 617 is further distilledto obtain a solution containing an enriched concentration offormaldehyde and/or formic acid. This enriched solution, in turn, iscontacted with the carbon-containing catalyst. As noted above, thisreduction treatment can occur in one or more catalyst holding tank(s) inreactor system 603.

At least another portion 649 of the adiabatic crystallizer overhead 617and/or at least a portion 651 of the withdrawn decantate 624 may bepurged (i.e., discharged) from the system as waste. In a continuoussystem, this purge helps to reduce the amount of impurity buildup in thesystem. This purged waste may, in turn, be further treated to removeimpurities by techniques known in the art, such as those described abovefor the purged waste streams of the centrifuge downstream of anon-adiabatic crystallizer. At least a portion 652 of the withdrawndecantate 624 may alternatively be sent to the non-adiabatic evaporativecrystallizer 663.

The primary N-(phosphonomethyl)glycine product slurry 625 withdrawn fromthe bottom of the adiabatic crystallizer 615 contains the bulk of theN-(phosphonomethyl)glycine product. The slurry 625 is typically passedthrough a centrifuge 655 to further concentrate the slurry 625 and forma wet cake 657 containing the N-(phosphonomethyl)glycine product.Normally, the concentration of the N-(phosphonomethyl)glycine product inthe wet cake 657 is at least about 95% (by weight of all compoundsbesides water).

At least a portion (preferably at least about 1%, more preferably fromabout 1% to about 67%, even more preferably from about 20% to about 50%,still even more preferably from about 30% to about 40%, and still yeteven more preferably about 33%) of the centrate 661 (i.e., primarymother liquor) from the centrifuge 655, on the other hand, is sent to aheat-driven evaporative crystallizer 663, which provides heat to thecentrate 661 to vaporize water and small molecule impurities to form theevaporative crystallizer overhead vapor stream 665. Much of theN-(phosphonomethyl)glycine product precipitates in the liquid medium667. This liquid medium 667 is withdrawn from the non-adiabaticevaporative crystallizer 663 and introduced into a hydroclone 669, whichforms a product-rich stream 673 enriched in precipitatedN-(phosphonomethyl)glycine product and a solids-depleted stream 671. Theproduct-rich stream 673 is introduced into a centrifuge 675 which formsa centrate 677 (which is further depleted in precipitatedN-(phosphonomethyl)glycine product) and an N-(phosphonomethyl)glycineproduct wet cake 679. In instances where the entire centrate 661 fromthe centrifuge 655 is not all introduced into the non-adiabaticcrystallizer 663, a portion 695 of the centrate 661 may be recycled backto the adiabatic crystallizer 615 and/or purged from the system viapurge stream 693 and treated using, for example, the various liquidwaste treatments discussed above.

In the embodiment shown in FIG. 13, at least a portion of theheat-driven evaporative crystallizer overhead 665 may be recycled backto the reactor system 603. Often, a portion 685 is recycled back to thereactor system 603 and used as a source of water for dissolving theN-(phosphonomethyl)iminodiacetic acid substrate to form the feed stream601 for the oxidation reaction zone(s). Particularly where the catalystis a carbon-containing catalyst, a portion of the heat-drivenevaporative crystallizer overhead 665 also may advantageously be used toreduce the catalyst surface. This is due to the fact that theevaporative crystallizer overhead 665 often contains formaldehyde and/orformic acid, which both act as reducing agents, particularly towardcarbon-containing catalysts. The reduction treatment may occur in one ormore catalyst holding tank(s) of the reactor system 603.

At least another portion of the heat driven evaporative crystallizeroverhead 665 is normally purged from the system as purge stream 683. Ina continuous system, this purge 683 helps to reduce the amount ofimpurity buildup (particularly small molecule impurity buildup) in thesystem. The purged waste 683 may, in turn, be further treated to removeimpurities, as discussed above for the purged overhead streams foradiabatic crystallizers and non-adiabatic crystallizers discussed above.

At least a portion 689 of the centrate 677 from the centrifuge 675 ispreferably recycled back to the heat-driven evaporative crystallizer 663(and/or to the adiabatic crystallizer 615) for further recovery of theN-(phosphonomethyl)glycine product. Alternatively (or in addition), aportion 691 of the centrate 677 is recycled back to the oxidationreaction zone(s) of the reactor system 603 to convert unreactedN-(phosphonomethyl)iminodiacetic acid substrate in the centrate 677 intoN-(phosphonomethyl)glycine product. A portion 687 of the centrate 677 isalso normally purged from the system. In a continuous system, this purge687 helps to reduce the amount of impurity buildup (particularly largermolecule impurity buildup) in the system. The purged waste 687 may befurther treated to remove impurities by, for example, the sametechniques described above for the liquid purged wastes discussed abovefor adiabatic and non-adiabatic crystallizers.

In an alternative embodiment, rather (or in addition to) feeding thecentrate 661 from the centrifuge 655 to the non-adiabatic crystallizer663, at least a portion 652 (preferably at least about 1%, morepreferably from about 1 to about 67%, even more preferably from about 20to about 50%, still even more preferably from about 30 to about 40%, andstill yet even more preferably about 33%) of the withdrawn decantate 624from the adiabatic crystallizer 615 is introduced into the non-adiabaticcrystallizer 663. In that instance, the centrate 661 may be, forexample, recycled to the adiabatic crystallizer 615 (via stream 695),recycled back to the reactor system 603, purged from the system (viastream 693), and/or introduced into the non-adiabatic crystallizer 663.

FIG. 14 shows an example an especially preferred embodiment wherein theprocess comprises an adiabatic crystallizer 319 and a non-adiabaticcrystallizer 343. Here, the adiabatic crystallizer 319 and non-adiabaticcrystallizer 343 operate in a semi-parallel manner.

Many of the various streams shown in FIG. 14 are analogous to thosedescribed above for non-adiabatic crystallizers and adiabaticcrystallizers alone. An aqueous feed stream 301 comprising anN-(phosphonomethyl)iminodiacetic acid substrate is introduced along withoxygen into an oxidation reactor system 303 comprising one or moreoxidation reaction zone(s), wherein the N-(phosphonomethyl)iminodiaceticacid substrate is oxidatively cleaved in the presence of a catalyst toform an oxidation reaction product solution 305. The oxidation reactionproduct solution 305 withdrawn from the last oxidation reaction zone ofreactor system 303 is then introduced into a pre-crystallizer flash tank306 to reduce the pressure and flash out much of the dissolved CO₂. Inthe embodiment shown in FIG. 14, the resulting liquid stream 308 isfiltered with a catalyst filter 307 to remove a heterogenous particulatecatalyst suspended in liquid stream 308 and form a catalyst recyclestream 309 and filtrate 311. The filtrate 311 is divided into pluralfractions and a portion 317 (i.e., a primary fraction of the reactionproduct solution) is introduced into the adiabatic crystallizer 319 toproduce a primary product slurry comprising precipitatedN-(phosphonomethyl)glycine product crystals and primary mother liquor,while another portion 315 (i.e., a secondary fraction of the reactionproduct solution) is introduced into the non-adiabatic heat-drivenevaporative crystallizer 343 to produce an evaporative crystallizationslurry 344 (i.e., a secondary product slurry) comprising precipitatedN-(phosphonomethyl)glycine product crystals and secondary mother liquor.In such an embodiment, the portion 315 of the filtrate 311 which isintroduced into the evaporative crystallizer 343 may first be introducedinto a crystallizer feed tank (not shown), where it is mixed with thesolids-depleted hydroclone stream 351 and/or recycled centrate 365. Inaddition to providing a location for mixing, such a feed tank alsoprovides a timing buffer to hold materials during, for example, startupand shut down of the process.

Operation of the adiabatic crystallizer 319 produces vapor 369 (i.e.,the adiabatic crystallizer overhead) discharged from the top of thecrystallizer, a decantate (i.e., primary mother liquor) stream 323withdrawn from the crystallizer and a primary crystallization productslurry 321 comprising precipitated crystallineN-(phosphonomethyl)glycine product and primary mother liquor removedfrom the bottom of the crystallizer. Preferably, at least a portion (andpreferably all) of the adiabatic crystallizer overhead 369 and/ordecantate 323 withdrawn from the adiabatic crystallizer 319 is/arerecycled back to the oxidation reaction zone(s) of the reactor system303. Typically, at least a portion 377 of the adiabatic crystallizeroverhead 369 and/or a portion 324 of the withdrawn decantate 323 is/arerecycled back to the reactor system 303 and used as a source of waterfor the oxidation reaction zone(s).

At least a portion 325 of the withdrawn decantate 323 may alternativelybe sent to the non-adiabatic evaporative crystallizer 343. At least aportion of the adiabatic crystallizer overhead 369 can be indirectlyrecycled (via stream 379) back to the reactor system 303 by being usedto reduce the catalyst surface. As noted above, this reduction treatmentoften occurs in a catalyst holding tank(s) 373.

Because the solids-depleted liquid streams from the adiabaticcrystallizer 319 (i.e., the decantate stream 323) and subsequentcentrifuge 331 (i.e., the centrate stream 335) typically have a lowerconcentration of impurities (particularly of larger molecule impurities)than the solids-depleted stream (i.e., the centrate) 357 of thecentrifuge 355 (e.g., a solid bowl centrifuge) downstream of theheat-driven evaporative crystallizer 343, it is normally more preferableto recycle back to the reactor system 303 the entire withdrawn decantate323 from the adiabatic crystallizer 319, and optionally the entiresolids-depleted stream 335 from the centrifuge 331 downstream of theadiabatic crystallizer 319, while using the solids-depleted stream(i.e., the centrate) 357 of the centrifuge 355 downstream of theheat-driven evaporative crystallizer 343 for purging larger moleculeimpurities from the system (via purge stream 361). Purging isadvantageous for continuous systems because it reduces the rate ofcontaminant buildup in the system, thus making recycle of thesolids-depleted streams (i.e., streams 323 and 335) from the adiabaticcrystallizer 319 more feasible. Like the purged waste streams discussedabove, this purged waste 361 may be treated by, for example, furthercrystallization. It may also be treated by the techniques described bySmith in U.S. Pat. No. 5,606,107. It may additionally be treated, forexample, by the technique described in detail by Morgenstern et al. inU.S. Pat. No. 6,005,140. It may further be treated by the techniquedescribed in detail by Grabiak et al. in U.S. Pat. No. 4,851,131.

The solids-depleted stream 335 from the centrifuge 331 downstream of theadiabatic crystallizer 319, on the other hand, is typically entirelyrecycled, for example, back to the adiabatic crystallizer 319 (viastream 337) or to the reactor system 303 (via stream 341). Thesolids-depleted stream 351 from the hydroclone 349 downstream of theheat-driven evaporative crystallizer 343 can be recycled back to theevaporative crystallizer. Any non-purged portion 365 of thesolids-depleted stream 357 from the centrifuge 355 downstream of theevaporative crystallizer 343 is typically recycled back to theevaporative crystallizer. At least a portion of the heat-drivenevaporative crystallizer overhead 345 is typically purged from thesystem via stream 347, although a portion 350 optionally may be recycledback to the oxidation reactor system 303 directly via stream 383 orindirectly via stream 381 to be used as a reducing agent for thecatalyst.

Preferably, from about 30% to about 85%, more preferably from about 50%to about 80%, and even more preferably from about 65% to about 75% ofthe oxidation reaction mixture absent the catalyst (i.e., stream 311) isintroduced into the adiabatic crystallizer 319 via stream 317 as theprimary fraction, while the remaining portion is introduced into thenon-adiabatic heat-driven crystallizer 343 via stream 315 as thesecondary fraction. The weight ratio of the secondary fraction 315 tothe N-(phosphonomethyl)iminodiacetic acid substrate fed into the systemis preferably from about 0.1 to about 9, more preferably from about 2 toabout 7, even more preferably from about 2.5 to 4.

Embodiments operating an adiabatic crystallizer and a heat-drivenevaporative crystallizer in a semi-parallel manner (such as the oneshown in FIG. 14) are typically more preferred than embodimentsoperating an adiabatic crystallizer and an evaporative crystallizer inseries. This stems, for example, from the fact that for a givenevaporative crystallizer size, greater crystallization capacitygenerally may be obtained where the crystallizers are in parallel. Thisprovides more flexibility in retrofitting existing plants.

The system in FIG. 14 produces two N-(phosphonomethyl)glycine productwet cake streams: a wet cake stream 333 from centrifuge 331 downstreamof the adiabatic crystallizer 319, and a wet cake stream 359 fromcentrifuge 355 downstream of the heat-driven evaporative crystallizer343. Normally, the wet cake stream 333 from the adiabatic crystallizer319 preferably has an N-(phosphonomethyl)glycine product concentrationof at least 90% (by weight of all compounds besides water), morepreferably at least 95% (by weight of all compounds besides water), andeven more preferably at least about 99% (by weight of all compoundsbesides water), while the wet cake stream 359 from the evaporativecrystallizer 343 has an N-(phosphonomethyl)glycine product concentrationof at least about 85% (by weight of all compounds besides water), morepreferably at least about 90% (by weight of all compounds besideswater), and even more preferably at least about 95% (by weight of allcompounds besides water). Typically, the purity of the wet cake 333 fromthe adiabatic crystallizer 319 section is greater than the purity of thewet cake 359 from the heat-driven evaporative crystallizer 343.

It is often advantageous to combine these wet cakes 333 and 359. Thiscombination allows lesser purity levels in the wet cake 359 from theevaporative crystallizer 343 to be tolerated due to the greater puritylevels normally obtained in the wet cake 333 issuing from the adiabaticcrystallizer 319. Thus, for example, if 45% of the combined wet cake isfrom the evaporative crystallizer 343 and 55% of the combined wet cakeis from the adiabatic crystallizer 319, the purity level of the wet cake359 from the evaporative crystallizer 343 may be as low as 90.2% byweight to achieve a combined purity level of at least 95% by weightwhere the wet cake 333 from the adiabatic crystallizer 319 is 99% byweight pure. Normally, it is desirable for the combined wet cake to havean N-(phosphonomethyl)glycine product concentration of at least about95% (by weight of all compounds besides water), and more preferablyabout 96% (by weight of all compounds besides water).

FIG. 14A shows an example of another especially preferred embodimentwherein the process comprises an adiabatic crystallizer and anon-adiabatic heat-driven evaporative crystallizer operated in asemi-parallel manner as described above in FIG. 14. However, thisreaction system further includes a secondary reactor system comprisingone or more secondary oxidation reaction zone(s) used in conjunctionwith the fraction of the reaction product mixture from the primaryoxidation reactor system sent to the heat-driven evaporativecrystallizer.

Many of the various streams shown in FIG. 14A are analogous to thosedescribed above for the reaction system shown in FIG. 14 in which theadiabatic crystallizer 319 and the heat-driven evaporative crystallizer343 are operated in a semi-parallel manner. An aqueous feed stream 301comprising an N-(phosphonomethyl)iminodiacetic acid substrate isintroduced along with oxygen into a primary oxidation reactor system 303comprising one or more oxidation reaction zone(s), wherein theN-(phosphonomethyl)iminodiacetic acid substrate is oxidatively cleavedin the presence of a catalyst to form a reaction product solution 305comprising N-(phosphonomethyl)glycine product and unreactedN-(phosphonomethyl)iminodiacetic acid substrate. The reaction productsolution 305 from the primary reactor system 303, after catalystseparation (e.g., by filtration) if necessary, is divided into pluralfractions comprising a primary fraction 317 and a secondary oxidationreactor feed fraction 315. The primary fraction 317 is introduced intothe adiabatic crystallizer 319 and is cooled by flash evaporation ofwater therefrom under reduced pressure conditions to recover theN-(phosphonomethyl)glycine product as described above. The secondaryoxidation reactor feed fraction 315 is introduced into a secondaryoxidation reactor system 316 comprising one or more oxidation reactionzones in which unreacted N-(phosphonomethyl)iminodiacetic acid substrateis oxidized to produce a secondary oxidation reactor effluent 318comprising N-(phosphonomethyl)glycine product. The secondary reactorfeed fraction 315 may be cooled prior to introduction into the secondaryoxidation reactor system 316 to remove exothermic heat generated in theprimary oxidation reactor system 303 and reduce production ofby-products. Reactor effluent 318 is introduced into the non-adiabaticheat-driven evaporative crystallizer wherein water is evaporatedtherefrom to precipitate and recover the N-(phosphonomethyl)glycineproduct as described above.

In the reaction system in FIG. 14A it should be understood that theprimary and secondary reactor systems 303 and 316, respectively, mayinclude one or more oxidation reaction zone(s) provided by variousreactor configurations including, for example, the continuous reactorsystems described hereinabove. For purposes of illustration, the primaryreactor system 303 may comprise a single stirred tank reactor as shownin FIGS. 2, 2A and 2B, two stirred tank reactors in series as shown inFIGS. 3-6 or one or more fixed bed reactors as shown in FIG. 8. In oneembodiment, the primary reactor system 303 comprises one or moreoxidation reaction zone(s) in series and the reaction product solution305 is withdrawn from the last oxidation reaction zone, in the series,and filtered as necessary. However, it should be understood that thereaction product solution 305 may be divided before the last oxidationreaction zone of the primary reactor system 303 such that the primaryfraction 317 passes through the remaining oxidation reaction zone(s) ofthe primary reactor system before being introduced into the adiabaticcrystallizer 319.

Preferably, the secondary reactor system 316 comprises one or moreoxidation reaction zone(s) provided by one or more fixed bed reactors orstirred tank reactors utilizing a particulate catalyst slurry orcombinations thereof. However, fixed bed reactors are generally morepreferred since a catalyst recycle mechanism including a catalyst filtercan be avoided in the secondary reactor system 316. Moreover, concernsregarding dissipation of exothermic reaction heat and temperaturecontrol that arise when a fixed bed reactor serves as the firstoxidation reaction zone in the primary reactor system 303 are largelycircumvented in the secondary reactor system 316 since the majority ofthe N-(phosphonomethyl)iminodiacetic acid substrate is preferablyoxidized in the primary reactor system 303. Accordingly, the oxidationreaction zone(s) within the secondary reactor system may be operatedadiabatically. In accordance with an especially preferred embodiment,the secondary reactor system 316 comprises a single oxidation reactionzone provided by a fixed bed reactor. Preferably, such a fixed bedreactor is operated with cocurrent gas and liquid flows through theoxidation reaction zone.

The presence of the secondary reactor system 316 in the reaction systemshown in FIG. 14A permits the primary reactor system 303 to beconfigured and operated in a manner which allows a higher concentrationof unreacted N-(phosphonomethyl)iminodiacetic acid substrate in thereaction mixture 305. For example, the second or subsequent oxidationreaction zone(s) of the primary reactor system 303 may be sizedconsiderably smaller or eliminated completely (i.e., the primary reactorsystem 303 may comprise a single oxidation reaction zone). However, theconcentration of unreacted N-(phosphonomethyl)iminodiacetic acidsubstrate in the reaction product solution 305, and thus in the primaryfraction 317 sent to the adiabatic crystallizer 319, is neverthelesspreferably maintained sufficiently low to avoid precipitation ofunreacted N-(phosphonomethyl)iminodiacetic acid substrate at theprevailing stream temperature. For typical operating temperatures of theadiabatic crystallizer (e.g., 60° C.), the concentration ofN-(phosphonomethyl)iminodiacetic acid substrate in the reaction productsolution 305 is no greater than about 2% by weight. However, in order totake advantage of the presence of the secondary reactor system 316 whichpermits the primary reactor system 303 to be sized and operated moreeconomically, the concentration of N-(phosphonomethyl)iminodiacetic acidsubstrate in the reaction product solution 305 is preferably at leastabout 0.2% by weight and more preferably at least about 0.5% by weight.

Unreacted N-(phosphonomethyl)iminodiacetic acid substrate in the primaryfraction 317 introduced into the adiabatic crystallizer 319 ispreferably recovered and returned to the primary reactor system 303 viathe decantate 323 withdrawn from the adiabatic crystallizer 319, as wellas by optionally recycling at least a portion of the solids-depletedstream 335 from the centrifuge 331 downstream of the adiabaticcrystallizer 319 to the primary reactor system via stream 341. Byemploying a high dewatering centrifuge (e.g., a vertical basketcentrifuge) as centrifuge 331, recovery of unreactedN-(phosphonomethyl)iminodiacetic acid substrate in the solids-depletedstream 335 is enhanced, while advantageously minimizing the fraction ofunreacted N-(phosphonomethyl)iminodiacetic acid substrate contained inwet cake stream 333.

In a further embodiment of the present invention, the system shown inFIG. 13 may be modified by the addition of a secondary reactor system tofurther oxidize unreacted N-(phosphonomethyl)iminodiacetic acidsubstrate in the centrate 661 from the centrifuge 655 upstream of thenon-adiabatic evaporative crystallizer 663.

Particularly where the N-(phosphonomethyl)glycine product isN-(phosphonomethyl)glycine itself, it has long been known that theproduct may be converted to various salts or esters to increase itssolubility in water so that it is more readily amenable to commercialuse. This is generally discussed, for example, by Franz in U.S. Pat.Nos. 3,977,860 and 4,405,531. Preferred salts ofN-(phosphonomethyl)glycine include, for example, alkali metal salts(particularly the potassium salt), alkanolamine salts (particularly themonoethanolamine salt), alkyl amine salts (particularly theisopropylamine salt), and alkyl sulfonium salts (particularly thetrimethyl sulfonium salt). The isopropylamine salt ofN-(phosphonomethyl)glycine is particularly preferred. See, e.g., Bugg etal., U.S. Pat. No. 5,994,269. This salt typically has a significantlygreater activity than the free acid, and is, for example, roughly 40times as soluble as the free acid at 25° C.).

In some embodiments of this invention, the N-(phosphonomethyl)glycineproduct formed in the oxidation reaction zone(s) comprises an ester orsalt which is sufficiently great to form a mixture having the desiredcommercial concentration. In those instances, the desirability forprocess steps (e.g., crystallization, hydrocycloning, centrifugation,and the like) for concentrating the product may be significantly reducedor entirely eliminated. This is especially true if the catalyst is thedeeply reduced catalyst discussed above, which typically forms areaction mixture having a low concentration of impurities andconsequently requiring little or no purification.

In some embodiments, for example, the N-(phosphonomethyl)glycine productformed in the oxidation reaction zone(s) is the isopropylamine salt ofN-(phosphonomethyl)glycine. At the more preferred oxidation operatingtemperatures (i.e., from about 95 to about 105° C.), such a product willremain soluble at concentrations of up to about 50% by weight orgreater. The salt product may be formed in the oxidation reactionzone(s) by (a) using the isopropylamine salt ofN-(phosphonomethyl)iminodiacetic acid as the substrate, (b) introducingisopropylamine into the oxidation reaction zone(s) to aminate theoxidation product while in the oxidation reaction zone(s), and/or (c)introducing isopropylamine into a vessel downstream of the oxidationreaction zone(s) and before catalyst filtration. Where there are morethan one oxidation reaction zones in series, it is normally preferable(although not absolutely necessary) to use the isopropylamine salt ofN-(phosphonomethyl)iminodiacetic acid as the substrate and/or introduceisopropylamine into the first of the oxidation reaction zones.Regardless, at least one equivalent (and more preferably slightly morethan one equivalent) of isopropylamine cations are preferably presentper mole of N-(phosphonomethyl)glycine product formed. It should berecognized these principles with respect to forming the isopropylaminesalts also generally apply to forming other salts, e.g., alkali metalsalts (particularly the potassium salt), alkanolamine salts(particularly the monoethanolamine salt), alkyl amine salts besides theisopropylamine salt, and alkyl sulfonium salts (particularly thetrimethyl sulfonium salt).

EXAMPLES

The following examples are simply intended to further illustrate andexplain the present invention. This invention, therefore, should not belimited to any of the details in these examples.

Example 1 Measuring Pore Volume of Carbon Support

A Micromeritics ASAP 2000 surface area and pore volume distributioninstrument was used to acquire the data. Total surface areadetermination involves exposing a known weight of a solid to somedefinite pressure of a non-specific adsorbate gas at a constanttemperature, e.g., at the temperature of liquid nitrogen, −196° C.During equilibration, gas molecules leave the bulk gas to adsorb ontothe surface which causes the average number of molecules in the bulk gasto decrease which, in turn, decreases the pressure. The relativepressure at equilibrium, p, as a fraction of the saturation vaporpressure, p_(o), of the gas is recorded. By combining this decrease inpressure with the volumes of the vessel and of the sample, the amount(i.e., the number of molecules) of gas adsorbed may be calculated byapplication of the ideal gas laws. These data are measured at relativepressures (p/p_(o)) of approximately 0.1 to 0.3 where the Brunauer,Emmett and Teller (BET) equation for multi-layer adsorption typicallyapplies. With the number of adsorbed gas molecules known, it is possibleto calculate the surface area using the “known” cross-sectional area ofthe adsorbate. For cases where only physical adsorption due to Van derWaals forces occurs (i.e., Type I Langmuir isotherms) the determinationof surface area from the observed changes in pressure is accomplishedusing the BET equation. Pore size and pore size distributions arecalculated by obtaining relative pressure data approaching p/p_(o)=1,i.e., in the regime where multi-layer adsorption and capillarycondensation occur. By applying the Kelvin equation and methodsdeveloped by Barrett, Joyner and Halenda (BJH), the pore volume and areamay be obtained.

Example 2 High-Temperature Deoxygenation of a Carbon Support

The high-temperature deoxygenation procedures described in the followingexamples may be used with any carbon support to produce a deoxygenatedcarbon support.

Single-Step High-Temperature Deoxygenation #1 Using NH₃/H₂O Gas

An activated carbon support (2.5 g) was placed into a 1.9 cm I.D.×40.6cm length quartz tube. The tube was connected to a gas stream resultingfrom sparging a 70 to 100 ml/min. N₂ stream through a 70° C., 10% NH₄OHaqueous solution. The quartz tube then was placed into a preheated 30.5cm tubular furnace and pyrolyzed at 930° C. for 60 min. and then cooledto room temperature under a dry N₂ atmosphere without contacting anyair.

Single-Step High-Temperature Deoxygenation #2 Using NH₃/H₂O Gas

An activated carbon support (3.55 g) was placed into a 1.9 cm I.D.×35.6cm long quartz tube. The tube was connected to streams of 50 ml/min. ofNH₃ gas and 89 ml/min. of steam and then placed into a preheated 30.5 cmtubular furnace and pyrolyzed at 930° C. for 30 minutes. The tubesubsequently was cooled to room temperature under a dry N₂ atmospherewithout any contact with air.

To show the advantages of deoxygenating the carbon support beforedispersing the noble metal onto the surface of the support, theperformances of the following two catalysts were compared: one having acarbon support, which was deoxygenated using the above treatment beforeplatinum was dispersed onto its surface; and one having an SA-30 carbonsupport (Westvaco Corp. Carbon, Department Covington, Va.) which wasused as received from Westvaco. Platinum was dispersed onto the surfacesof the carbon supports using the technique described in Example 3 below.The catalysts then were reduced. In one experiment, the catalysts werereduced using NaBH₄ (See Example 12 for protocol). In a secondexperiment, the catalysts were reduced by heating them in 20% H₂ and 80%argon for 8 hours at 640° C.

The reduced catalysts were used to catalyze the oxidation ofN-(phosphonomethyl)iminodiacetic acid to N-(phosphonomethyl)glycine(i.e., “glyphosate”) using the reaction conditions set forth in Example5. Table 1 shows the results. Use of the deoxygenated carbon supportresulted in smaller CO desorption values, less noble metal leaching,higher formaldehyde activity, and shorter reaction times.

TABLE 1 Effect of Deoxygenating the Carbon Support before DispersingNoble Metal onto Its Surface CO desorption Deoxygenation from carbon Ptin soln. (μg/g CH₂O (mg/g glyph. Reaction time¹ treatment support(mmole/g) Reduction glyph. prod.) prod.) (min.) Single-step 0.23 NaBH₄8.6 28.5 35.1 high-temperature Reduced deoxygenation #2 (Ex. 12) SA-30,used as received 1.99 same 54.3 43.1 62.7 Single-step 0.23 8 hrs at 4.815.6 29.8 high-temperature 640° C. in deoxygenation #2 20% H2, 80% ArSA-30, used as received 1.99 same 31 19.7 50.7 ¹When ≧98% of theN-(phosphonomethyl)iminodiacetic acid has been consumed.

Example 3 Depositing Platinum onto the Surface of a Carbon Support

Twenty grams of NUCHAR activated carbon SA-30 (Westvaco Corp., CarbonDepartment, Covington, Va.) was slurried in 2 L of water for 2 hours.Then, 2.81 grams of H₂PtCl₆ dissolved in about 900 ml of water was addeddropwise over a period of 3 to 4 hours. After the H₂PtCl₆ solution wascompletely added, the slurry was stirred for 90 more minutes. The pH ofthe slurry then was readjusted to 10.5 using NaOH, and stirred for 10 to14 more hours. The resulting slurry was filtered and washed with wateruntil the filtrate reached a constant conductivity. The wet cake wasdried at 125° C. under vacuum for 10 to 24 hours. This material produced5% platinum on carbon upon reduction.

It should be recognized that the above procedure may be used to depositplatinum onto the surface of other carbon supports as well.

Example 4 High-Temperature Hydrogen Reduction of a Carbon Support

Approximately 5.8 g of a dried, unreduced catalyst consisting of 5%platinum on a NUCHAR SA-30 carbon support (Westvaco Corp., CarbonDepartment, Covington, Va.) was dehydrated in-situ at 135° C. in argonfor one hour before being reduced at 640° C. with 20% H₂ in argon for 11hours. Upon cooling to room temperature under 20% H₂ in argon, thecatalyst was ready to use.

It should be recognized that the above procedure may be used to heatother carbon supports as well.

Example 5 Use of the Catalyst to OxidizeN-(phosphonomethyl)iminodiacetic acid to N-(Phosphonomethyl)glycine

This example demonstrates the use of high-temperature gas-phasereduction to improve catalyst performance.

An Aldrich catalyst consisting of 5% platinum on an activated carbonsupport (catalog No. 20,593-1, Aldrich Chemical Co., Inc., Milwaukee,Wis.) was heated at 640° C. for 4-6 hours in the presence of 20% H₂ and80% argon. Subsequently, it was used to catalyze the oxidation ofN-(phosphonomethyl)iminodiacetic acid to Glyphosate. Its performance wascompared to the performance of a sample of the Aldrich catalyst whichwas used as received from Aldrich.

The N-(phosphonomethyl)iminodiacetic acid oxidation reaction wasconducted in a 200 ml glass reactor using 11.48 g ofN-(phosphonomethyl)iminodiacetic acid, 0.5% catalyst (dry basis), atotal reaction mass of 140 g, a temperature of 90° C., a pressure of 50psig, a stir rate of 900 rpm, and an oxygen flow rate of 100 ml/min.

Table 2 shows the results. The high-temperature hydrogen-reducedcatalyst had less leaching, better formaldehyde activity, and producedless N-methyl-N-(phosphonomethyl)glycine. Also, reaction time wasshortened by 30% when the high-temperature hydrogen-reduced catalyst wasused.

TABLE 2 N-(phosphonomethyl)iminodiacetic acid Oxidation Results for 5%Pt on Activated Carbon (Aldrich Cat. No. 20,593-1) Catalyst As ReceivedHigh-Temp., H₂ Reduced NPMIDA (%) 0.4619 0.4430N-(phosphonomethyl)glycine 5.58 5.54 (%) HCO₂H (mg/g glyph. prod.) 46.9935.87 CH₂O (mg/g glyph. prod.) 32.96 14.60 NMG (mg/g glyph. prod.) 3.581.32 AMPA (ppm) 172.5 182.0 End Point (min.) 64.67 44.17 Pt in soln.(μg/g glyph. prod.) 32.26 10.50 % of Pt Lost 0.72 0.232

Example 6 Further Examples Showing Use of Catalyst to OxidizeN-(phosphonomethyl)iminodiacetic acid to N-(Phosphonomethyl)glycine

This example demonstrates using the high-temperature, gas-phasereduction treatment and ammonia washing to improve catalyst performance.

The performances of six catalysts in catalyzing theN-(phosphonomethyl)iminodiacetic acid oxidation were compared. Thesecatalysts were: (a) a catalyst consisting of 5% platinum on an activatedcarbon support (Catalog No. 33,015-9, Aldrich Chemical Co., Inc.,Milwaukee, Wis.); (b) the catalyst after being washed with ammonia(ammonia washing was conducted using the same technique described inExample 10 except that the pH of the catalyst slurry was adjusted to andmaintained at 11.0 rather than 9.5); (c) the catalyst after being heatedat 75° C. in 20% H₂ and 80% argon for 4-6 hours (GPR@75° C.); (d) thecatalyst after being heated at 640° C. for 4-6 hours in the presence of20% H₂ and 80% argon (GPR@640° C.); and (e) two catalysts after beingwashed with ammonia and then heated at 640° C. for 4-6 hours in thepresence of 20% H₂ and 80% argon. The N-(phosphonomethyl)iminodiaceticacid oxidation reaction conditions were the same as in Example 5.

Table 3 shows the results. The untreated catalyst showed relatively highleaching and poor formaldehyde activity. High-temperature gas-phasereduction at 640° C. in the presence of H₂ leads to the greatestdecrease in leaching and increase in formaldehyde activity. Heating thecatalyst at 75° C. in 20% H₂ at 75° C. decreased leaching to a lesserextent, but did not enhance the formaldehyde activity.

TABLE 3 NPMIDA Oxidation Results for 5% Pt on Activated Carbon (AldrichCat. No. 33,015-9) NH₃ wash NH₃ wash + GPR NH₃ wash + GPR CatalystAs-received w/o GPR¹ GPR@75° C. GPR@640° C. @640° C. @640° C. NPMIDA (%)ND ND ND 0.097 0.083 ND Glyphosate (%) 5.87 5.65 5.81 5.89 5.85 5.91HCO₂H (mg/g glyph. prod.) 43.46 43.65 38.97 42.14 46.91 52.12 CH₂O (mg/gglyph. prod.) 19.39 22.73 19.85 13.78 15.70 17.61 NMG (mg/g glyph.prod.) 1.27 0.89 0.89 1.00 1.31 1.68 AMPA (ppm) 149.4 147.6 134.6 349.8324.8 283.8 End Point (min.) 39.33 44.33 38 31.42 34.33 33.33 Pt insoln. (μg/g glyph. 42.59 40.71 27.54 5.26 5.30 4.23 prod.) % of Pt Lost1 0.92 0.64 0.12 0.12 0.1 ¹“GPR” means reduction in H₂ ²“ND” means nonedetected.

In the next experiment, five catalysts were analyzed while catalyzingthe N-(phosphonomethyl)iminodiacetic acid oxidation. These catalystswere: (a) a catalyst consisting of 5% platinum on NUCHAR SA-30 (WestvacoCorp., Carbon Department, Covington, Va.); (b) the catalyst after beingtreated with NaBH₄ (see Example 12 for protocol); (c) the catalyst afterbeing heated at 75° C. in 20% H₂ and 80% argon for 4-6 hours (GPR@75°C.); (d) the catalyst after being heated at 640° C. in 20% H₂ and 80%argon for 4-6 hours (GPR@640° C.); (e) the catalyst after being washedwith ammonia (using the same technique described in Example 10) and thenheated at 640° C. in 20% H₂ and 80% argon for 4-6 hours. The reactionconditions were the same as those in Example 5.

Table 4 shows the results. The untreated catalyst showed relatively highplatinum leaching and low formaldehyde activity. The catalyst alsoshowed high leaching and low formaldehyde activity after being treatedwith NaBH₄, as did GPR@75° C. In contrast, GPR@640° C. showed a greaterformaldehyde activity and less leaching.

TABLE 4 NPMIDA Oxidation Results Using 5% Pt on NUCHAR SA-30 NH₃ wash +GPR Catalyst Unreduced NaBH₄ red. GPR@75° C. GPR@640° C. @640° C.Glyphosate (%) 2.50 5.71 4.92 5.17 5.19 HCO₂H (mg/g glyph. prod.) 59.5651.14 57.85 30.85 38.21 CH₂O (mg/g glyph. prod.) 115.28 43.13 48.5219.67 20.79 NMG (mg/g glyph. prod.) 1.64 2.17 6.41 0.37 1.73 AMPA (ppm)58.16 193.9 174.0 138.5 156.3 End point (min.) 62.67 62.67 70.67 50.6759.33 Pt in soln. (μg/g glyph. prod.) 84.00 54.29 81.30 30.95 19.27 % ofPt Lost 0.84 1.24 1.6 0.64 0.4

Example 7 Effect of C/O and O/Pt Ratios at the Surface of the Catalyst

The carbon atom to oxygen atom ratio and the oxygen atom to platinumatom ratio at the surfaces of various fresh catalysts were analyzedusing a PHI Quantum 2000 ESCA Microprobe Spectrometer (PhysicalElectronics, Eden Prairie, Minn.). The surface analysis was performed byelectron spectroscopy for chemical analysis (“ESCA”) with the instrumentin a retardation mode with the analyzer at fixed band pass energy(constant resolution). The analysis entails irradiation of the samplewith soft X-rays, e.g., Al K_(α) (1486.6 eV), whose energy is sufficientto ionize core and valence electrons. The ejected electrons leave thesample with a kinetic energy that equals the difference between theexciting radiation and the “binding energy” of the electron (ignoringwork function effects). Because only the elastic electrons, i.e., thosethat have not undergone energy loss by any inelastic event, are measuredin the photoelectron peak, and because the inelastic mean free path ofelectrons in solids is short, ESCA is inherently a surface sensitivetechnique. The kinetic energy of the electrons is measured using anelectrostatic analyzer and the number of electrons are determined usingan electron multiplier. The data are presented as the number ofelectrons detected versus the binding energy of the electrons. ESCAsurvey spectra were taken using monochromatic Al K_(α) x-rays forexcitation of the photoelectrons with the analyzer set for a 117 eV bandpass energy. The X-ray source was operated at 40 watts power and datawere collected from the 200 μm spot on the sample being irradiated.These conditions give high sensitivity but low energy resolution. Thespectra were accumulated taking a 1.0 eV step size across the regionfrom 1100 eV to 0 eV and co-adding repetitive scans to achieveacceptable signal/noise in the data. The elements present wereidentified and quantified using the standard data processing andanalysis procedures provided with the instrumentation by the vendor.From the relative intensities of the photoelectron peaks, the relativeatomic concentrations of the elements Pt/C/O are obtained. ESCA analysisis generally cited as having a precision of ±20% using tabulatedresponse factors for a particular instrument configuration.

Table 5 shows the C/O and O/Pt ratios at the surface of each freshcatalyst, and the amount of leaching for each of the catalysts during asingle-cycle N-(phosphonomethyl)iminodiacetic acid oxidation reaction.

TABLE 5 Effects of C/O and O/Pt Ratios During NPMIDA Oxidation¹Reduction Treatment Pt After Depositing C/O O/Pt in Soln. CH₂O CatalystNoble Metal Ratio Ratio (μg/g)² (mg/g)³ 5% Pt on NaBH₄ 23.7 3 ND⁴deoxygenated Reduced carbon⁵ same Pt(II)⁶ 35.3 17 1.2 24.44 640° C./9hr/10% H₂ same NaBH₄ Reduced 21.1 3 6.9 Aldrich Cat. 640° C./6 hr/20% H₂67.9 3 5.2 13.78 No. 33015-9 same 75° C./6 hr/20% H₂ 13.4 10 27.5 19.85same Used as Received 13.3 10 42.6 19.39 Aldrich Cat. 640° C./6 hr/20%H₂ 45.2 7 10.5 21.90 #20593-1 NH₃ wash/pH = 11 same 640° C./6 hr/20% H₂37.7 10 10.5 14.60 same Used as Received 9.1 26 32.3 32.96 5% Pt on 640°C./7 hr/20% H₂ 67.7 8 19.3 20.79 SA-30 NH₃ wash/pH = 9.5 Westvaco carbonsame 640° C./8 hr/20% H₂ 63.3 8 30.9 19.67 same 75° C./7 hr/20% H₂ 13.232 81.3 48.52 ¹The reaction conditions were the same as those used inExample 5. ²μg Pt which leached into solution per gram Glyphosateproduced. ³mg formaldehyde per gram Glyphosate produced. ⁴“ND” meansnone detected. ⁵Carbon support deoxygenated with the single-stephigh-temperature technique #2 of Example 2. ⁶Pt deposited usingdiamminedinitrito P(II) as described in Example 11.

Example 8 Analysis of Catalyst Surface Using Thermogravimetric Analysiswith In-Line Mass Spectroscopy (TGA-MS)

The concentration of oxygen-containing functional groups at the surfacesof various fresh catalysts was determined by thermogravimetric analysiswith in-line mass spectroscopy (TGA-MS) under helium. To perform thisanalysis, a dried sample (100 mg) of fresh catalyst is placed into aceramic cup on a Mettler balance. The atmosphere surrounding the samplethen is purged with helium using a flow rate 150 ml/min. at roomtemperature for 10 minutes. The temperature subsequently is raised at10° C. per minute from 20 to 900° C., and then held at 900° C. for 30minutes. The desorptions of carbon monoxide and carbon dioxide aremeasured by an in-line mass spectrometer. The mass spectrometer iscalibrated in a separate experiment using a sample of calcium oxalatemonohydrate under the same conditions.

Table 6 shows the amount of carbon monoxide desorbed per gram of eachcatalyst using TGA-MS, and the amount of leaching for each of thecatalysts during a single-cycle N-(phosphonomethyl)iminodiacetic acidoxidation reaction using the same reaction conditions as in Example 5.As Table 6 shows, leaching tends to decrease as the amount of COdesorption decreases, and is particularly low when the desorption is nogreater than 1.2 mmole/g (mmole CO desorbed per gram of catalyst).

TABLE 6 Effects of Oxygen-Containing Functional Groups Which Desorb fromCatalyst Surface as CO during TGA-MS Reduction TGA-MS Pt in Soln. CH₂OCatalyst Treatment (mmole/g)¹ (μg/g)² (mg/g)³ Aldrich Cat. 640° C./6hr/20% H₂ 0.41 5.2 13.78 #33015-9 same 640° C./6 hr/20% H₂ 0.38 5.315.70 NH₃ wash/pH = 9.5 same 75° C./6 hr/20% H₂ 1.87 27.5 19.85 same NH₃wash/pH = 9.5 1.59 40.7 22.73 same Used as Received 1.84 42.6 19.39¹mmole of CO per gram of catalyst ²μg of noble metal which leaches intosolution per gram of Glyphosate produced ³mg of formaldehyde per gram ofGlyphosate produced

Example 9 Effect of Temperature During High-Temperature Gas-PhaseReduction

This example demonstrates the effects of using various temperatures whenheating the catalyst in the presence of a reducing agent.

An unreduced catalyst having 5% platinum on an activated carbon support(which was deoxygenated using the single-step high-temperaturedeoxygenation technique #2 described in Example 2 before the platinum isdeposited) was heated at various temperatures in 10% H₂ and 90% argonfor about 2 hours. The catalyst then was used to catalyze theN-(phosphonomethyl)iminodiacetic acid oxidation reaction. The reactionwas conducted in a 250 ml glass reactor using 5 gN-(phosphonomethyl)iminodiacetic acid, 0.157% catalyst (dry basis), 200g total reaction mass, a temperature of 80° C., a pressure of 0 psig,and an oxygen flow rate of 150 ml/min.

The results are shown in Table 7. Increasing the reduction temperaturefrom 125° C. to 600° C. reduces the amount of noble metal leaching andincreases the formaldehyde oxidation activity during the oxidationreaction of N-(phosphonomethyl)iminodiacetic acid into Glyphosate.

TABLE 7 Effects of Reduction Temperature Reduction Temperature Pt inSoln. CH₂O C/O O/Pt (° C.) (normalized¹) (normalized²) Ratio Ratio 1251.00 0.41 26 13 200 0.44 0.80 27 14 400 0.18 0.93 42 10 500 0.14 0.95 3214 600 0.06 1.00 40 11 ¹A normalized value of 1.00 corresponds to thehighest amount of Pt observed in solution during this experiment. ²Anormalized value of 1.00 corresponds to the highest formaldehydeactivity during this experiment.

Example 10 Washing the Catalyst with Ammonia

An unreduced catalyst (6.22 g) consisting of 5% platinum on an activatedcarbon support (which was deoxygenated using the single-stephigh-temperature deoxygenation technique #2 described in Example 2before the platinum was deposited onto the support) was slurried in 500ml of water for 30 minutes. Afterward, the pH of the slurry was adjustedto 9.5 with diluted aqueous ammonia, and the slurry was stirred for onehour, with aqueous ammonia being periodically added to maintain the pHat 9.5. The resulting slurry was filtered and washed once with about 300ml of water. The wet cake then was dried at 125° C. under vacuum forabout 12 hours. This catalyst was heated at 640° C. for 11 hours in 10%H₂ and 90% argon, and then compared with two other catalysts consistingof 5% platinum on NUCHAR activated carbon: (a) one reduced at roomtemperature with NaBH₄ (see Example 12 for protocol), and (b) one heatedat 640° C. in 10% H₂ and 90% argon for 11 hours. The reactions were thesame as those in Example 5.

The results are shown in Table 8. Platinum leaching was the lowest withthe catalyst which was washed with ammonia before high-temperaturehydrogen reduction.

TABLE 8 Effects of Ammonia Washing CH₂O HCO₂H NMG Pt in soln. Catalyst(mg/g)¹ (mg/g) (mg/g) (μg/g) NH₃-washed, 10.62 28.79 0.83 0.50High-Temp., H₂-reduced High-temp., 14.97 27.82 1.38 4.64 H₂-reducedRoom-Temp., 28.51 70.16 2.59 8.64 NaBH₄-reduced ¹These quantities areper gram Glyphosate produced.

Example 11 Use of a Less Oxidizing Noble Metal Precursor

Platinum was deposited on an activated carbon support usingdiamminedinitrito platinum (II). Approximately 20 g of an activatedcarbon support was deoxygenated using the single-step high-temperaturedeoxygenation technique #2 described in Example 2. Next, it was slurriedin 2 L of water for 2 hours. Approximately 51.3 g of a 3.4% solution ofdiamminedinitrito platinum (II), diluted to 400 g with water, then wasadded dropwise over a period of 3-4 hours. After addition was complete,stirring was continued for 90 more minutes. The pH was re-adjusted to10.5 by adding diluted aqueous NaOH, and stirring was conducted for10-14 more hours. The slurry then was filtered and washed with aplentiful amount of water until the filtrate reached constantconductivity. The wet cake was dried at 125° C. under vacuum for 10-24hours. The resulting catalyst was heated at 640° C. for 4-6 hours in 10%H₂ and 90% argon.

A control was prepared using H₂PtCl₆ to deposit platinum onto the samecarbon. The control was heated under the same conditions as the catalystprepared using diamminedinitrito platinum (II).

These catalysts were compared while catalyzing theN-(phosphonomethyl)iminodiacetic acid oxidation reaction. The reactionconditions were the same as those in Example 5.

The catalyst prepared using diamminedinitrito platinum (II) showed lessleaching than the control. Only 1.21 μg platinum per gram of Glyphosateproduced leached into solution, which was about three times better thanthe control.

Example 12 Reducing the Catalyst Surface Using NaBH₄

The purpose of this example is to demonstrate the effects of reducingthe catalyst using NaBH₄.

Approximately 5 g of an activated carbon support (which was deoxygenatedusing the single-step high-temperature deoxygenation technique #2described in Example 2 before the platinum was deposited onto thesupport) was slurried with 85 ml of distilled water in a 250 ml roundbottom flask. The slurry was stirred in a vacuum for about 1 hour. Next,0.706 g of H₂PtCl₆ in 28 ml of distilled water was added to the slurryat a rate of about 1 ml per 100 seconds with the vacuum still beingapplied. After stirring overnight in the vacuum, the reactor was broughtto atmospheric pressure by admitting a flow of N₂. After allowing theslurry to settle, approximately 30 ml of colorless supernatant wasdecanted. The remaining slurry was transferred to a 100 ml Teflon roundbottom. At this point, the pH was adjusted to 12.2 with 0.3 g of NaOH.Then, 2.3 ml of NaBH₄ in 14 M NaOH was added at 0.075 ml/min.Subsequently, the resulting slurry was stirred for one hour, filtered,and washed five times with 50 ml of distilled water. The catalyst thenwas dried at 125° C. and 6 mmHg for 12 hours.

The resulting catalyst was used to catalyze theN-(phosphonomethyl)iminodiacetic acid oxidation. The reaction wasconducted in a 300 ml stainless steel reactor using 0.5% catalyst, 8.2%N-(phosphonomethyl)iminodiacetic acid, a total reaction mass of 180 g, apressure of 65 psig, a temperature of 90° C., an agitation rate of 900rpm, and an oxygen feed rate of 72 ml/min.

A control experiment also was conducted at the same reaction conditionsusing 5.23% platinum on an activated carbon support (which wasdeoxygenated using the single-step high-temperature deoxygenationtechnique #2 described in Example 2 before the platinum was depositedonto the support).

Table 9 shows the results using the NaBH₄-reduced catalyst, and Table 10shows the results of the control experiment. Reducing with NaBH₄ reducedthe amount of noble metal leaching. It also reduced the amount offormaldehyde and NMG after a period of use.

TABLE 9 Results Using Catalyst Treated with NaBH₄ Run # 1 2 3 4 5 6Glyphosate (%) 5.79 5.81 5.75 5.74 5.79 5.77 NPMIDA (%) 0.23 0.08 0.130.22 0.13 0.13 CH₂O (mg/g glyph) 28.5 31.5 47.8 38.8 41.6 45.8 HCO₂H(mg/g glyph) 70.2 90.5 100.5 96.6 98.8 99.0 AMPA/MAMPA (%) 0.02 0.010.01 0.01 0.01 0.01 NMG (mg/g glyph) 2.6 3.6 3.6 4.2 4.7 4.7 Pt in Soln.8.64 8.60 5.22 6.96 6.91 5.20 (μg/g glyph.) % of Pt Lost 0.20 0.20 0.120.16 0.16 0.12

TABLE 10 Results Using Catalyst which was not treated with NaBH₄ Run # 12 3 4 5 6 Glyphosate (%) 5.36 5.63 5.37 5.50 5.56 5.59 NPMIDA (%) 0.180.15 0.25 0.21 0.18 0.23 CH₂O (%) 20.9 23.6 38.4 44.2 47.7 58.3 HCO₂H(%) 27.8 63.8 96.5 98.4 102.2 102.0 AMPA/MAMPA (%) 0.04 0.02 0.04 0.020.02 0.03 NMG (mg/g glyph) 1.5 3.0 5.4 6.9 10.6 7.3 Pt in Soln 63.6 62.244.7 34.6 28.8 28.6 (μg/g glyph.) % of Pt Lost 1.30 1.34 0.92 0.73 0.610.61

Example 13 Use of Bismuth as a Catalyst-Surface Promoter

A 500 g solution was prepared consisting of 10⁻³ M Bi(NO₃)₃.5H₂O in 10⁻³M formic acid solution. This solution was added to 500 g of a 5%formaldehyde solution containing 6.0 g of 5% platinum on an activatedcarbon support. The solution was stirred at 40° C. under N₂ overnightand then filtered with a Buchner funnel. An aliquot was dried andsubsequently analyzed by X-ray fluorescence. The catalyst had a loss ondrying (“LOD”) of 63%. The dry catalyst was found to containapproximately 3% bismuth and 4% platinum.

The following were placed into a 300 ml stainless steel autoclave: 16.4g of N-(phosphonomethyl)iminodiacetic acid; 4.16 g of an activatedcarbon catalyst, 0.68 g of the above catalyst consisting of 3%bismuth/4% platinum on its surface, and 179.4 g of water. The reactionwas conducted at a pressure of 65 psig, a temperature of 90° C., anoxygen flow rate of 38 ml/min., and a stir rate of 900 rpm. The reactionwas allowed to proceed until the N-(phosphonomethyl)iminodiacetic acidwas depleted. The product solution was separated from the catalyst viafiltration and the solution was neutralized with 6 g of 50% NaOHsolution. The catalyst was recycled with no purge through 5 runs.Analysis of the product solution was done for each run. Two controlsalso were conducted in the same manner as above except that the 0.68 gof the Bi/Pt/carbon catalyst was omitted.

The results are shown in Table 11. The runs having the Bi/Pt/carboncatalyst produced lower levels of formaldehyde, formic acid, and NMG inthe product.

TABLE 11 NPMIDA Oxidation Results Using Pt/Bi/C Catalyst CONTROL CONTROL#1 #2 1ST RUN 2ND RUN 3RD RUN 4TH RUN 5TH RUN Glyphosate (%) 5.7 5.595.69 5.72 5.87 5.74 5.68 NPMIDA (%) ND ND 0.04 0.07 0.085 0.04 0.046AMPA (%) 0.034 0.031 0.015 0.009 0.008 DBNQ¹ DBNQ CH₂O (mg/g glyph.prod.) 142 138 28 31 34 38 42 HCO₂H (mg/g glyph. 56 57 DBNQ 7 14 17 23prod.) AMPA/MAMPA (%) 0.047 0.041 0.021 0.014 0.013 0.014 0.013 NMG(mg/g glyph. prod.) 16.3 19.3 0.7 0.9 1.4 2.3 2.6 ¹DBNQ = detectable,but not quantified

Example 14 Depositing a Tin Promoter on a Carbon Support

An activated carbon (20 g) was slurried in about 2 L of water. Next,0.39 g of SnCl₂.2H₂O was dissolved in 500 g of 0.5% HNO₃. The solutionwas added dropwise to the carbon slurry. After all the solution wasadded, the slurry was stirred for 2 hours. The pH then was adjusted to9.5, and the slurry was stirred for a few more hours. Next, the slurrywas filtered and washed with a plentiful amount of water until thefiltrate reached a constant conductivity. The wet cake was dried at 125°C. under vacuum to give 1% tin on carbon. Following drying, the 1% tinon carbon was calcined in argon at 500° C. for 6 hours.

To deposit platinum onto the carbon support, 5 g of the 1% tin on carbonfirst was slurried in about 500 ml of water. Then 0.705 g of H₂PtCl₆ wasdissolved in about 125 ml of water and added dropwise. After all theH₂PtCl₆ solution was added, the slurry was stirred for 2.5 hours. The pHthen was adjusted to 9.5 with diluted NaOH and stirring was continuedfor a few more hours. The slurry then was filtered and washed with aplentiful amount of water until the filtrate reached constantconductivity. The wet cake was dried at 125° C. under vacuum.

This technique produced a catalyst comprising 5% platinum and 1% tin oncarbon.

Example 15 Depositing an Iron Promoter onto a Carbon Support

Approximately 5 g of activated carbon was slurried in about 500 ml ofwater. Next, 0.25 g of FeCl₃.6H₂O was dissolved in 75 ml of water. Thesolution was added dropwise to the carbon slurry. After all the solutionwas added, the slurry was stirred for two hours. The slurry then wasfiltered and washed with a plentiful amount of water until the filtratereached a constant conductivity. The wet cake was dried at 125° C. undervacuum to give 1% iron on carbon. Following drying, the 1% iron oncarbon was calcined in argon at about 500° C. for 8 hours.

To deposit platinum onto the surface of the carbon support, 2.5 g of the1% iron on carbon first was slurried in about 180 ml of water. Then,0.355 g of H₂PtCl₆ was dissolved in about 70 ml of water and addeddropwise. After all the solution was added, the slurry was stirred forthree more hours. The pH then was adjusted to about 10.0 with dilutedNaOH and stirring was continued for a few more hours. Next, the slurrywas filtered and washed with a plentiful amount of water until thefiltrate reached a constant conductivity. The wet cake was dried at 125°C. under vacuum.

This technique produces a catalyst comprising 5% platinum and 1% iron oncarbon.

Example 16 Effect of Presence of Noble Metal on the Surface of theCarbon Support

This example shows the advantages of using a carbon support having anoble metal on its surface for effecting the oxidation ofN-(phosphonomethyl)iminodiacetic acid rather than a carbon-only catalysthaving no noble metal on its surface.

The N-(phosphonomethyl)iminodiacetic acid oxidation reaction wasconducted in the presence of a carbon-only catalyst which wasdeoxygenated using the single-step high-temperature deoxygenationtechnique #2 described in Example 2. The reaction was carried out in a300 ml stainless steel reactor using 0.365% catalyst, 8.2%N-(phosphonomethyl)iminodiacetic acid, a total reaction mass of 200 g, apressure of 65 psig, a temperature of 90° C., an agitation rate of 900rpm, and an oxygen feed rate of 38 ml/min.

Table 12 shows the reaction times (i.e., the time for at least 98% ofthe N-(phosphonomethyl)iminodiacetic acid to be consumed) of 5 cyclesfor the carbon-only catalyst. Table 12 also shows the reaction times forthe two Pt-on-carbon catalysts in Example 12 over 6 cycles under thereaction conditions described Example 12. As may be seen from Table 12,the deactivation of the carbon-only catalyst per cycle generally tendsto be greater (i.e., the reaction times tend to increase more per cycle)than the deactivation of the carbon catalysts which had a noble metal ontheir surfaces. The deactivation particularly appears to be less wherethe catalyst has been reduced with NaBH₄ after the noble metal wasdeposited onto the surface. Without being bound by any particulartheory, it is believed that the deactivation of the catalyst reducedwith NaBH₄ was less than the deactivation of the other Pt-on-carboncatalyst because the platinum on the NaBH₄ catalyst leached less thanthe platinum on the other Pt-on-carbon catalyst. See Example 12, Tables9 & 10.

TABLE 12 Results Using Catalyst which was not treated with NaBH₄ Run # 12 3 4 5 6 Run Time for 45.4 55.0 64.4 69.8 75.0 Carbon-Only Catalyst(min.) Run Time for 35.1 NA¹ NA 35.2 35.8 35.8 5% platinum on CarbonCatalyst which was Reduced w/ NaBH₄ (min.) Run Time for 40.4 42.0 44.244.1 44.9 52.7 5.23% platinum on Carbon Catalyst (min.) ¹Not availabledue to temperature problems.

Example 17 The Effect of Using a Catalyst Comprising a Noble MetalAlloyed with a Catalyst-Surface Promoter

This example shows the advantages of a catalyst comprising platinumalloyed with iron.

1. Catalyst Comprising Platinum Alloyed with Iron

To prepare the catalyst comprising platinum alloyed with iron,approximately 10 grams of an activated carbon was slurried in about 180ml of water. Next, 0.27 grams of FeCl₃.6H₂O and 1.39 grams of H₂PtCl₆hydrate were co-dissolved in about 60 ml of water. This solution wasadded dropwise to the carbon slurry over a period of about 30 minutes.During the addition, the pH of the slurry dropped and was maintained atfrom about 4.4 to about 4.8 using a dilute NaOH solution (i.e., a 1.0 to2.5 molar solution of NaOH). Afterward, the slurry was stirred for 30more minutes at a pH of about 4.7. The slurry then was heated under N₂to 70° C. at a rate of about 2° C./min. while maintaining the pH atabout 4.7. Upon reaching 70° C., the pH was raised slowly over a periodof about 30 minutes to 6.0 with addition of the dilute NaOH solution.The stirring was continued for a period of about 10 min. until the pHbecame steady at about 6.0. The slurry was then cooled under N₂ to about35° C. Subsequently, the slurry was filtered, and the cake was washedwith approximately 800 ml of water 3 times. The cake was then dried at125° C. under a vacuum. This produced a catalyst containing 5 wt. %platinum and 0.5 wt. % iron on carbon upon heating at 690° C. in 20% H₂and 80% Ar for 1-6 hr.

This catalyst was analyzed via electron microscopy, as described in moredetail in Example 19. An image obtained through TEM of the carbonsupport showed that the alloyed metal particles were highly dispersedand uniformly distributed throughout the carbon support (the white dotsrepresent the metal particles; and the variations in the backgroundintensity are believed to represent the change of the local density ofthe porous carbon). The average size of the particles was about 3.5 nm,and the average distance between particles was about 20 nm. A highenergy resolution X-ray spectra from an individual metal particle of thecatalyst showed that both platinum and iron peaks were present (thecopper peaks originated from the scattering of the copper grids).Quantitative analysis of the high energy resolution X-ray spectra fromdifferent individual metal particles showed that the composition of theparticles, within experimental error, did not vary with the size or thelocation of the metal particles on the catalyst surface.

2. Catalyst in which Platinum was Less Alloyed with Iron

To prepare the Pt/Fe/C catalyst in which the platinum was less alloyedwith iron (i.e., this catalyst has less platinum alloyed with iron thandoes the first catalyst described in this example), the platinum andiron were deposited sequentially onto the surface of the carbon support.Approximately 5 grams of an activated carbon was slurried in about 500ml of water. The pH was adjusted to about 5.0 with 1N HCl. Next, about0.25 grams of FeCl₃.6H₂O was dissolved in 75 ml of water. This solutionwas added dropwise to the carbon slurry over a period of about 60 min.After all the solution was added, the slurry was stirred for about 2hours. The pH was adjusted to 9.5 with the dilute NaOH solution, and theslurry was stirred for a few more hours. Afterward, the slurry wasfiltered and washed with a plentiful amount of water. The wet cake wasdried at 125° C. under vacuum to produce 1 wt. % iron on carbon.Following drying, this 1 wt. % iron on carbon was reduced with anatmosphere containing 20% H₂ and 80% Ar at 635° C. for 1-6 hr. About 2.5grams of this 1 wt. % iron on carbon was slurried in 250 ml of water.Next, about 0.36 grams of H₂PtCl₆ hydrate was dissolved in 65 ml ofwater, which, in turn, was added dropwise to the slurry over a period ofabout 60 min. After all the solution was added, the slurry was stirredfor 2 hours. The slurry then was filtered and washed with a plentifulamount of water. The cake was then re-slurried in 450 ml of water. Afteradjusting the pH of the slurry to 9.5 with the dilute NaOH solution, theslurry was stirred for about 45 min. Next, the slurry was filtered andwashed once with 450 ml of water. The wet cake was the dried at 125° C.under vacuum. This produced a catalyst containing 5 wt. % platinum and 1wt. % iron on carbon upon reduction by heating to a temperature of 660°C. in an atmosphere containing 20% H₂ and 80% Ar for 1-6 hr.

3. Comparison of the Two Catalysts

These two catalysts were compared while catalyzing theN-(phosphonomethyl)iminodiacetic acid oxidation reaction. The reactionconditions were the same as those in Example 5. Table 13 shows theresults. The first catalyst described in this example (i.e., thecatalyst comprising a greater amount of platinum alloyed with iron) hadgreater stability with respect to CH₂O & HCO₂H activities; the secondcatalyst described in this example (i.e., the catalyst comprising alower amount of platinum alloyed with iron) deactivated rapidly. Inaddition, the first catalyst retained almost half of its iron contentover 25 cycles, while the second catalyst lost most of its iron in thefirst cycle.

TABLE 13 Comparison of Catalyst Having Pt/Fe Alloy with Catalyst HavingLess Pt/Fe Alloy cycle 1 cycle 2 cycle 3 cycle 4 cycle 5 cycle 6 cycle 7cycle 8 cycle 9 cycle 10 cycle 11 cycle 12 cycle 13 Alloyed Pt & Fe CH₂O(mg/g glyph. 10.49 9.23 6.04 4.92 4.44 5.08 5.24 prod.) HCO₂H (mg/gglyph. 19.91 29.64 27.84 25.62 27.99 29.73 28.95 prod.) NMG (mg/g glyph.0.22 0.44 0.28 0 0 0 0 prod.) Pt in soln. 5.08 4.87 3.6 3.06 (μg/gglyph. prod.) % of Fe Lost 44 1.9 1.2 0.8 Less alloyed Pt & Fe CH₂O(mg/g glyph. 10.16 10.7 12.24 13.56 14.68 prod.) HCO₂H (mg/g glyph.27.23 37.72 45.01 54.57 61.14 prod.) NMG (mg/g glyph. 0 0.98 1.23 1.77 2prod.) Pt in soln. 3.83 3.36 3.54 3.44 3.32 (μg/g glyph. prod.) % of FeLost 86 3.2 1.4 1.8 1.4

Example 18 Preparation of a Pt/Fe/Sn on Carbon Catalyst

Approximately 10 grams of an activated carbon was slurried in about 90ml of water. Next, about 0.2 g of SnCl₂.2H₂O was dissolved in 250 ml of0.025 M HCl. The solution was added dropwise to the carbon slurry. Afterall the solution was added, the slurry was stirred for 3 hr. The pH thenwas slowly adjusted to 9.0 with a diluted NaOH solution (i.e., a 1.0 to2.5 molar solution of NaOH), and the slurry was stirred for a few morehours. Next, the slurry was filtered and washed with a plentiful amountof water until the filtrate reached a constant conductivity. The wetcake was dried at 125° C. under vacuum. This produced 0.9 wt. % tin oncarbon. About 6 grams of this 0.9 wt. % tin on carbon was slurried inabout 500 ml of water. Then approximately 0.23 grams of Fe(NO₃)₃.9H₂Oand 0.85 grams of H₂PtCl₆ were co-dissolved in about 150 ml of water andadded dropwise to the slurry. After all the solution was added, theslurry was stirred for 4 hours, and then filtered to remove excess iron(˜80 wt. %). The wet cake was re-slurried in 480 ml of water. After thepH of the slurry was adjusted to 9-10 with the dilute NaOH solution, theslurry was stirred for a few more hours. Next, the slurry was filteredand washed with a plentiful amount of water until the filtrate reached aconstant conductivity. The wet cake was dried at 125° C. under vacuum.This produced a catalyst containing 4.9 wt. % Pt, 0.9 wt. % tin and 0.1wt. % iron on carbon upon high-temperature reduction by heating at700-750° C. in 20% H₂ and 80% Ar for 1-6 hr.

Example 19 Electron Microscopy Characterization of Catalysts

Electron microscopy techniques were used to analyze the size, spatialdistribution, and composition of the metal particles of catalystsprepared in Example 17. Before analyzing the catalyst, the catalyst wasfirst embedded in an EM Bed 812 resin (Electron Microscopy Sciences,Fort Washington, Pa.). The resin was then polymerized at about 60° C.for approximately 24 hr. The resulting cured block was ultramicrotomedinto slices having a thickness of about 50 nm. These slices were thentransferred to 200 mesh copper grids for electron microscopyobservation.

High-resolution analytical electron microscopy experiments were carriedout in a Vacuum Generators dedicated scanning transmission electronmicroscope (model no. VG HB501, Vacuum Generators, East Brinstead,Sussex, England) with an image resolution of less than 0.3 nm. Themicroscope was operated at 100 kV. The vacuum in the specimen chamberarea was below about 10⁻⁶ Pa. A digital image acquisition system (ESVision Data Acquisition System, EmiSpec Sys., Inc., Tempe, Ariz.) wasused to obtain high-resolution electron microscopy images. A windowlessenergy dispersive X-ray spectrometer (Link LZ-5 EDS Windowless Detector,Model E5863, High Wycombe, Bucks, England) was used to acquire highenergy resolution X-ray spectra from individual metal particles. Becauseof its high atomic-number sensitivity, high-angle annular dark-field(HAADF) microscopy was used to observe the metal particles. An electronprobe size of less than about 0.5 nm was used to obtain the HAADFimages, and a probe size of less than about 1 nm was used to obtain highenergy resolution X-ray spectra.

Example 20 Effect of a Supplemental Promoter

This example shows the use and advantages of mixing a supplementalpromoter with a carbon-supported, noble-metal-containing oxidationcatalyst.

A. Comparison of Effects on a NPMIDA Oxidation Reaction Caused by Mixinga Carbon-Supported, Noble-Metal-Containing Catalyst with Various Amountsand Sources of Bismuth

Several single batch N-(phosphonomethyl)iminodiacetic acid oxidationreactions were conducted. In each reaction, a different source and adifferent amount of bismuth were added to the reaction medium. Thesource of bismuth was either (BiO)₂CO₃, Bi(NO₃)₃.5H₂O, or Bi₂O₃. Theamount of bismuth used corresponded to a bismuth toN-(phosphonomethyl)iminodiacetic acid mass ratio of 1:10,000; 1:2,000;or 1:1,000. A control was also conducted wherein no bismuth was added.

Each N-(phosphonomethyl)iminodiacetic acid oxidation reaction wasconducted in the presence of a catalyst containing 5% by weight platinumand 0.5% by weight iron (this catalyst was prepared using a methodsimilar to that described in Example 17). The reaction was carried outin a 1000 ml stainless steel reactor (Autoclave Engineers, Pittsburgh,Pa.) using 2.5 g catalyst (0.5% by weight of the total reaction mass),60.5 g N-(phosphonomethyl)iminodiacetic acid (12.1% by weight of thetotal reaction mass), 1000 ppm formaldehyde, 5000 ppm formic acid, atotal reaction mass of 500 g, a pressure of 110 psig, a temperature of100° C., and an agitation rate of 1000 rpm. The oxygen feed rate for thefirst 22 minutes was 392 ml/min., and then 125 ml/min. until theN-(phosphonomethyl)iminodiacetic acid was essentially depleted.

Table 14 shows the results. In all the runs where a bismuth compound wasadded, the formaldehyde, formic acid, and NMG levels were less thanthose observed in the control.

TABLE 14 Direct Addition of Various Sources and Amounts of Bismuth AMPA/Amt. & source Glyph. NPMIDA CH₂O HCO₂H MAMPA NMG Run Time of Bi Added(%)** (%)** (mg/g)*** (mg/g)*** (mg/g)*** (mg/g)*** (min.) 0 (control)8.2 ND 4.0 22.5 9.4 2.0 39.3 0.0074 g 8.1 ND 2.6 3.8 10.9 ND 54.1(BiO)₂CO₃ (100 ppm*) 0.037 g 7.8 ND 1.8 1.4 14.5 ND 58.2 (BiO)₂CO₃ (500ppm) 0.074 g 7.7 ND 2.0 1.3 16.4 ND 60.2 (BiO)₂CO₃ (1000 ppm) 0.0141 g8.1 ND 2.4 3.0 11.2 ND 53.2 Bi(NO₃)₃•5H₂O (100 ppm) 0.070 g 7.7 ND 1.91.4 14.4 ND 58.5 Bi(NO₃)₃•5H₂O (500 ppm) 0.141 g 7.6 ND 2.0 1.2 16.2 ND59.2 Bi(NO₃)₃•5H₂O (1000 ppm) 0.0067 g 8.1 ND 2.5 3.5 13.9 ND 48 Bi₂O₃(100 ppm) 0.034 g 7.6 ND 2.0 1.4 15.1 ND 58.7 Bi₂O₃ (500 ppm) 0.067 g7.6 ND 2.0 1.2 17.3 ND 60.6 Bi₂O₃ (1000 ppm) *ppm means a ratio of Bi toN-(phosphonomethyl)iminodiacetic acid equaling 1:1,000,000 **(mass ÷total reaction mass) × 100% ***mg ÷ grams of glyphosate produced “ND”means none detectedB. Effect of Bismuth Addition on Subsequent NPMIDA Oxidation BatchesContacted with the Catalyst

Four 6-run experiments (i.e., during each of the 4 experiments, 6 batchreactions were conducted in sequence) were conducted to determine theeffect of (1) the initial bismuth addition on reaction runs subsequentto the initial bismuth addition, and (2) adding additional bismuth inone or more of the subsequent reaction runs.

All 4 experiments were conducted using a catalyst containing 5% byweight platinum and 0.5% by weight iron (this catalyst was preparedusing a method similar to that described in Example 17). During each6-run experiment, the same catalyst was used in each of the 6 runs(i.e., after the end of a run, the reaction product solution wasseparated and removed from the catalyst, and a new batch ofN-(phosphonomethyl)iminodiacetic acid was then combined with thecatalyst to begin a new run). The reaction was carried out in a 1000 mlstainless steel reactor (Autoclave Engineers) using 2.5 g catalyst (0.5%by weight of the total reaction mass), 60.5 gN-(phosphonomethyl)iminodiacetic acid (12.1% by weight of the totalreaction mass), 1000 ppm formaldehyde, 5000 ppm formic acid, a totalreaction mass of 500 g, a pressure of 110 psig, a temperature of 100°C., and an agitation rate of 1000 rpm. The oxygen feed rate for thefirst 22 minutes was 392 ml/min., and then 125 ml/min. until theN-(phosphonomethyl)iminodiacetic acid was essentially depleted.

In the control experiment, no bismuth was introduced into the reactionzone during any of the 6 runs. In the three other experiments, 0.034grams of bismuth(III) oxide (i.e., Bi₂O₃) were introduced into thereaction medium at the beginning of the first reaction run. In one ofthese experiments, the bismuth oxide was only introduced into thereaction zone at the beginning of the first reaction run. In anotherexperiment, 0.034 g of bismuth(III) oxide was introduced into thereaction medium at the beginning of the first and fourth reaction runs.In the final experiment, 0.034 g of bismuth(III) oxide was introducedinto the reaction medium at the beginning of all 6 reaction runs.

Tables 15, 16, 17, and 18 show the results. The one-time addition of thebismuth oxide (data shown in Table 16) tended to give the samebeneficial effects as adding the bismuth oxide every three runs (datashown in Table 17) or even every run (data shown in Table 18).

TABLE 15 Control Experiment: 6-Run NPMIDA Oxidation Reaction with NoBismuth Addition Sample (unless otherwise indicated, taken after approx.all NPMIDA consumed) Run 1 Run 2 Run 3 Run 4 Run 5 Run 6 Glyphosate 8.28.4 8.4 8.5 8.5 8.4 (%)* NPMIDA (%)* ND 0.006 0.008 ND ND ND CH₂O 3.12.4 2.0 2.6 3.2 3.8 (mg/g)** HCO₂H 16 23 22 25 30 40 (mg/g)** AMPA/ 7.56.9 6.3 5.5 5.8 5.9 MAMPA (mg/g)** NMG (mg/g)** 0.5 1.7 1.4 1.6 2.8 4.9Time (min.) 48.5 43.5 54.5 52.8 54.1 51.7 *(mass ÷ total reaction mass)× 100% **mg ÷ grams of glyphosate produced “ND” means none detected

TABLE 16 6-Run NPMIDA Oxidation Reaction with Bismuth Addition atBeginning of First Run Sample (unless otherwise indicated, taken afterapprox. all NPMIDA consumed) Run 1 Run 2 Run 3 Run 4 Run 5 Run 6Glyphosate 7.8 8.6 8.5 8.6 8.6 7.7 (%)* NPMIDA (%)* ND ND ND ND ND 0.005CH₂O 2.4 2.7 2.1 2.6 3.1 3.9 (mg/g)** HCO₂H DBNQ DBNQ DBNQ DBNQ DBNQDBNQ (mg/g)** AMPA/ 15 11 10 9.9 8.6 10 MAMPA (mg/g)** NMG (mg/g)** NDND ND ND ND ND Time (min.) 60.1 62.4 64.1 62.6 66.9 62 *(mass ÷ totalreaction mass) × 100% **mg ÷ grams of glyphosate produced “ND” meansnone detected “DBNQ” means detected, but not quantified

TABLE 17 6-Run NPMIDA Oxidation Reaction with Bismuth Addition atBeginning of 1st and 4th Runs Sample (unless otherwise indicated, takenafter approx. all NPMIDA consumed) Run 1 Run 2 Run 3 Run 4 Run 5 Run 6Glyphosate 7.8 8.4 8.5 8.5 8.5 8.6 (%)* NPMIDA (%)* ND ND ND ND ND NDCH₂O 2.3 2.6 2.6 3.2 3.6 3.5 (mg/g)** HCO₂H 3.4 3.1 3.2 2.9 3.3 3.5(mg/g)** AMPA/ 14 11 10 11 9.3 8.9 MAMPA (mg/g)** NMG (mg/g)** ND ND NDND ND ND Time (min.) 57.4 63.2 64.3 64.9 66 64.5 *(mass ÷ total reactionmass) × 100% **mg ÷ grams of glyphosate produced “ND” means nonedetected

TABLE 18 6-Run NPMIDA Oxidation Reaction with Bismuth Addition atBeginning of Every Run Sample (unless otherwise indicated, taken afterapprox. all NPMIDA consumed) Run 1 Run 2 Run 3 Run 4 Run 5 Run 6Glyphosate 7.8 8.5 8.2 8.3 8.3 8.3 (%)* NPMIDA (%)* ND ND ND ND ND NDCH₂O 2.4 2.8 3.2 2.9 3.4 4.0 (mg/g)** HCO₂H ND ND ND ND ND ND (mg/g)**AMPA/ 14 12 11 12 10 9.7 MAMPA (mg/g)** NMG (mg/g)** ND ND ND ND ND NDTime (min.) 56.4 62.4 64.8 62.8 66 66.1 *(mass ÷ total reaction mass) ×100% **mg ÷ grams of glyphosate produced “ND” means none detectedC. Effect of a One-Time Bismuth Addition Over 20 NPMIDA Oxidation RunsUsing a Platinum/Iron/Carbon Catalyst

Two 20-run experiments were conducted to determine the effect of aone-time bismuth addition on 20 N-(phosphonomethyl)iminodiacetic acidoxidation reaction runs.

Both experiments were conducted using a catalyst containing 5% by weightplatinum and 0.5% by weight iron (this catalyst was prepared using asimilar method to the method described in Example 17). During eachexperiment, the same catalyst was used in each of the 20 runs. Thereaction was carried out in a 1000 ml stainless steel reactor (AutoclaveEngineers) using 2.5 g catalyst (0.5% by weight of the total reactionmass), 60.5 g N-(phosphonomethyl)iminodiacetic acid (12.1% by weight ofthe total reaction mass), 1000 ppm formaldehyde, 5000 ppm formic acid, atotal reaction mass of 500 g, a pressure of 110 psig, a temperature of100° C., and an agitation rate of 1000 rpm. The oxygen feed rate for thefirst 22 minutes was 392 ml/min., and then 125 ml/min. until theN-(phosphonomethyl)iminodiacetic acid was essentially depleted. In thecontrol experiment, no bismuth was introduced into the reaction zoneduring any of the 20 runs. In the other experiment, 0.034 grams ofbismuth(III) oxide was introduced into the reaction medium at thebeginning of the first reaction run.

FIG. 15 compares the resulting formic acid concentration profiles. Theone-time introduction of bismuth into the reaction zone decreased theformic acid concentration over all 20 runs.

D. Effect of a One-Time Bismuth Addition Over 30 NPMIDA Oxidation RunsUsing a Platinum/Tin/Carbon Catalyst

Two 30-run experiments were conducted to determine the effect of aone-time bismuth addition on 30 N-(phosphonomethyl)iminodiacetic acidoxidation reaction runs.

Both experiments were conducted using a catalyst containing 5% by weightplatinum and 1% by weight tin (this catalyst was prepared using a methodsimilar to that described in Example 18). During each experiment, thesame catalyst was used in each of the 30 runs. Each run was carried outin a 300 ml reactor (made of alloy metal, Hastelloy C, AutoclaveEngineers) using 1.35 g catalyst (0.75% by weight of the total reactionmass), 21.8 g N-(phosphonomethyl)iminodiacetic acid (12.1% by weight ofthe total reaction mass), 1000 ppm formaldehyde, 5000 ppm formic acid, atotal reaction mass of 180 g, a pressure of 90 psig, a temperature of100° C., and an agitation rate of 900 rpm. The oxygen feed rate for thefirst 26 minutes was 141 ml/min., and then 45 ml/min. until theN-(phosphonomethyl)iminodiacetic acid was essentially depleted. In thecontrol experiment, no bismuth was introduced into the reaction zoneduring any of the 30 runs. In the other experiment, 0.012 grams ofbismuth (III) oxide was introduced into the reaction medium at thebeginning of the first reaction run.

FIG. 16 compares the resulting formic acid concentration profiles, FIG.17 compares the resulting formaldehyde concentration profiles, and FIG.18 compares the resulting NMG concentration profiles. Even after 30runs, the one-time introduction of bismuth into the reaction zonedecreased the formic acid concentration by 98%, the formaldehydeconcentration by 50%, and the NMG concentration by 90%.

E. Effect of Adding Bismuth to a Pt/Fe/C Catalyst that was PreviouslyUsed in 132 Batch NPMIDA Oxidation Reactions

A 14-run experiment was conducted to determine the effect mixing bismuthwith a used Pt/Fe/C catalyst. Before this experiment, the catalyst hadbeen used to catalyze 129 batch N-(phosphonomethyl)iminodiacetic acidoxidation reactions. The fresh catalyst (i.e., the catalyst before itwas used in the previous 129 N-(phosphonomethyl)iminodiacetic acidoxidation runs) was prepared using a method similar to the methoddescribed in Example 17, and contained 5% by weight platinum and 0.5% byweight iron.

The 14 N-(phosphonomethyl)iminodiacetic acid oxidation reaction runswere carried out in a 300 ml reactor (made of alloy metal, Hastelloy C,Autoclave Engineers) using 0.9 g of spent catalyst (0.5% by weight),21.8 g N-(phosphonomethyl)iminodiacetic acid (12.1% by weight), 1000 ppmformaldehyde, 5000 ppm formic acid, a total reaction mass of 180 g, apressure of 90 psig, a temperature of 100° C., and an agitation rate of900 rpm. The oxygen feed rate for the first 26 minutes was 141 mL/min.,and then 45 mL/min. until the N-(phosphonomethyl)iminodiacetic acid wasessentially depleted. At the beginning of the 4th run, 0.012 grams ofbismuth(III) oxide was introduced into the reaction zone.

FIG. 19 shows the effects that the bismuth addition at the 4th run hadon the formic acid, formaldehyde, and NMG by-product production.

F. Effect of Adding Bismuth to a Pt/Sn/C Catalyst that was PreviouslyUsed in 30 Batch NPMIDA Oxidation Reactions

An 11-run experiment was conducted to determine the effect of mixingbismuth with a used Pt/Sn/C catalyst. The catalyst had previously beenused to catalyze 30 batch N-(phosphonomethyl)iminodiacetic acidoxidation reactions. The fresh catalyst (i.e., the catalyst before itwas used in the previous 30 N-(phosphonomethyl)iminodiacetic acidoxidation runs) was prepared using a method similar to that described inExample 18, and contained 5% by weight platinum and 1% by weight tin.

The 11 N-(phosphonomethyl)iminodiacetic acid oxidation reaction runswere carried out in a 300 ml reactor (made of alloy metal, Hastelloy C,Autoclave Engineers) using 1.35 g of used catalyst (0.75% by weight ofthe total reaction mass), 21.8 g N-(phosphonomethyl)iminodiacetic acid(12.1% by weight of the total reaction mass), 1000 ppm formaldehyde,5000 ppm formic acid, a total reaction mass of 180 g, a pressure of 90psig, a temperature of 100° C., and an agitation rate of 900 rpm. Theoxygen feed rate for the first 26 minutes was 141 ml/min., and then 45mL/min. until the N-(phosphonomethyl)iminodiacetic acid was essentiallydepleted. At the beginning of the 4th run, 0.012 grams of bismuth(III)oxide was introduced into the reaction zone.

FIG. 20 shows the effects that the bismuth addition at the 4th run hadon the formic acid, formaldehyde, and NMG by-product production.

G. Effect of Bismuth Addition on Over 100 Subsequent NPMIDA OxidationBatches Contacted with the Catalyst

Two 125-run experiments were conducted to determine the effect ofbismuth addition on over 100 subsequent reaction runs using the samecatalyst.

Both experiments were conducted using a catalyst containing 5% by weightplatinum and 1% by weight tin (this catalyst was prepared using a methodsimilar to that described in Example 18). During each experiment, thesame catalyst was used in all the runs. The reaction was carried out ina stirred tank reactor using 0.75% catalyst (by weight of the totalreaction mass), 12.1% N-(phosphonomethyl)iminodiacetic acid (by weightof the total reaction mass), a pressure of 128 psig, and a temperatureof 100° C. The oxygen feed rate for the first part of each batchreaction (the exact amount of time varied with each batch from 14.9 to20.3 minutes, with times closer to 14.9 minutes being used for theearlier batches, and times closer to 20.3 minutes being used for thelater batches) was 1.3 mg/min. per gram total reaction mass, and then0.35 mg/min. per gram total reaction mass until theN-(phosphonomethyl)iminodiacetic acid was essentially depleted. Aportion of the reaction product from each batch was evaporated off andreturned to the reactor as a source of formaldehyde and formic acid toact as sacrificial reducing agents in the next batch reaction. Theamounts of formaldehyde and formic acid recycled back to the reactorranged from 100 to 330 ppm, and from 0 ppm to 2300 ppm (0 to 200 ppmformic acid after 25 batches following the addition of bismuth(III)oxide), respectively.

In the control experiment, no bismuth was introduced into the reactionzone during any of the 125 runs. In the other experiment, the catalystwas first used to catalyze 17 batches ofN-(phosphonomethyl)iminodiacetic acid. After catalyzing the 17th batch,the catalyst was substantially separated from the reaction product, andthe resulting catalyst mixture was transferred to a catalyst holdingtank where 9.0 mg of bismuth(III) oxide per gram of catalyst wereintroduced into the catalyst mixture. The catalyst was then used tocatalyze the oxidation of 107 subsequent batches ofN-(phosphonomethyl)iminodiacetic acid.

FIG. 21 compares the resulting formic acid concentration profiles, FIG.22 compares the resulting formaldehyde concentration profiles, and FIG.23 compares the resulting NMG concentration profiles. Even after 107runs, the one-time introduction of bismuth into a mixture with thecatalyst decreased the formic acid and NMG concentrations by roughly90%.

Example 21 Continuous Oxidation of NPMIDA to Glyphosate with PartiallySpent Catalyst and the Use of a Supplemental Promoter

This example demonstrates the continuous oxidation ofN-(phosphonomethyl) iminodiacetic acid (“NPMIDA”) toN-(phosphonomethyl)glycine (“glyphosate”) in a continuous oxidationreactor system utilizing a previously used catalyst and a supplementalpromoter. The experiment was designed to simulate the conditions thatmight prevail in a first reaction zone, particularly where crystallizermother liquor containing reaction product is recycled to the reactionzone.

The reaction was performed in a continuous reactor system utilizing a2-liter Hastelloy C autoclave (Autoclave Engineers Inc., Pittsburgh,Pa.). The reactor was equipped with an agitator having a 1.25″ diametersix-blade turbine impeller, which was operated at 1600 RPM. The liquidlevel in the reactor was monitored using a Drexelbrook Universal III™Smart Level™, with a teflon-coated sensing element. An internal coolingcoil was utilized to control the temperature within the reactor duringthe course of the reaction.

During operation, the reactor was continuously fed an aqueous slurryfeed material containing NPMIDA and a gaseous stream of oxygen. Theoxygen was introduced into the reaction medium through a frit locatednear the impeller. A liquid product stream containing the productN-(phosphonomethyl)glycine (“glyphosate”) was continuously withdrawnfrom the reactor through a frit, which allowed any catalyst charged tothe reactor to remain in the reaction medium. The product gas stream(containing CO₂ and unreacted oxygen) was continuously vented from thereactor headspace.

The reaction was begun by charging an aqueous slurry feed material (1420grams) and catalyst (29.3 grams or about 2 wt % catalyst in reactionmass) to the reactor. The aqueous slurry feed material contained NPMIDA(7.53% by weight), glyphosate (2.87% by weight), formaldehyde (2127 ppmby weight) and formic acid (3896 ppm by weight). The feed slurry alsocontained NaCl (about 450 ppm by weight) to mimic low level chlorideimpurity typically present in commercially available NPMIDA. Thecatalyst, which was prepared by a method similar to that described inExample 17 above, comprised platinum (5% by weight) and iron (0.5% byweight) on a particulate carbon support. The catalyst had beenpreviously used under conditions similar to those described in Example20.

The reactor was sealed to prevent any liquid inlet or outlet flow. Thereaction mixture was heated to about 105° C. and brought to a pressureof about 100 psig with nitrogen. Oxygen flow (1000 sccm) was initiatedand the reaction was run with no liquid inlet or outlet flow for about15 minutes. After this initial 15 minutes, slurry feed (70.4 g/min) wasinitiated, and reaction liquid was continuously withdrawn to maintain aconstant reactor level as indicated by the Drexelbrook level indicatordescribed above. After about 55 minutes, the oxygen flow was loweredslightly to 800 sccm. After about 280 minutes of operation at an oxygenflow of 800 sccm, Bi₂O₃ (0.0336 grams) was injected into the reactor asa supplemental promoter. The liquid product was analyzed with HPLC.Analytical results of the continuous oxidation reaction are shown inTable 19 below. Also, FIG. 24 shows profiles for formaldehyde and formicacid in the product liquid while the oxygen flow was 800 sccm.

TABLE 19 Oxidation Results from HPLC for Example 21 Time NPMIDAGlyphosate Formaldehyde Formic Acid (min)¹ (wt %) (wt %) (ppm) (ppm) 550.85 7.74 3977 4758 172 1.43 7.48 3078 5338 270 1.37 7.52 3137 5545 3472.41 6.87 2872 1395 405 2.42 6.97 2801 1385 464 2.48 6.99 2887 1474 4922.27 7.01 2881 1472 ¹Time after slurry feed started.

Example 22 Continuous Oxidation of NPMIDA to Glyphosate in the Presenceof a Previously Used Pt/Fe/C Catalyst

This example demonstrates the continuous oxidation of NPMIDA toglyphosate using a previously used heterogeneous particulate catalyst.The experiment was designed to simulate the conditions that mightprevail in a first reaction zone of a continuous reactor system,particularly where crystallizer mother liquor containing reactionproduct is recycled to the reaction zone.

The experiment was conducted in a continuous reactor system similar tothat described in Example 21 above. The reaction was begun by chargingan aqueous slurry feed material (1424 grams) and a heterogeneousparticulate catalyst (29.3 grams or about 2% catalyst by weight ofreaction mass) to the reactor. The aqueous slurry feed materialcontained NPMIDA (7.01% by weight), glyphosate (2.88% by weight),formaldehyde (2099.9 ppm by weight) and formic acid (4690 ppm byweight). The slurry feed also contained NaCl (about 450 ppm by weight)to mimic low level chloride impurity typically present in commerciallyavailable NPMIDA. The catalyst was prepared by a method similar to thatdescribed in Example 17 above and comprised platinum (5% by weight) andiron (0.5% by weight) on a particulate carbon support. The catalyst hadbeen previously used under conditions similar to those described inExample 20.

The reactor was sealed to prevent any liquid inlet or outlet flow. Thereaction mixture was heated to about 107° C. and brought to a pressureof about 100 psig with nitrogen. Oxygen flow (900 sccm) was initiatedand the reaction was run with no liquid inlet or outlet flow for about13 minutes. After this initial 13 minutes, slurry feed (70.4 g/min) wasinitiated, and reaction liquid was continuously withdrawn to maintain aconstant reactor level as indicated by the Drexelbrook level indicatordescribed in Example 21 above. The liquid product was analyzed withHPLC. Analytical results for the continuous oxidation reaction are shownin Table 20 below. Profiles for glyphosate produced and NPMIDA reactantremaining in the product liquid are shown in FIG. 25.

TABLE 20 Oxidation Results for Example 22 Time NPMIDA Glyphosate (min)¹(wt %) (wt %) 94 0.67 7.10 138 0.55 7.02 192 0.50 7.12 274 0.46 7.09 3580.47 7.06 ¹Time after slurry feed started.

Example 23 Continuous Oxidation of NPMIDA to Glyphosate in the Presenceof a Pt/Fe/C Catalyst

This example demonstrates the continuous oxidation of NPMIDA toglyphosate in the presence of a fresh Pt/Fe/C heterogeneous particulatecatalyst (at a relatively low catalyst concentration) over an extendedperiod of time. The experiment was designed to simulate the conditionsthat might prevail in a first reaction zone, particularly wherecrystallizer mother liquor containing reaction product is recycled tothe reaction zone.

The experiment was conducted in a continuous reactor system similar tothat described in Example 21 above. The reaction was begun by chargingthe reactor with an aqueous slurry feed material (1447 grams) and aheterogeneous particulate catalyst (3.63 grams or about 0.25 wt %catalyst in reaction mass). The aqueous slurry feed material containedNPMIDA (3.45% by weight), glyphosate (1.55% by weight), formaldehyde(1140 ppm by weight) and formic acid (2142 ppm by weight). The feedslurry also contained NaCl (about 450 ppm) to mimic low level chlorideimpurity typically present in commercially available NPMIDA. Thecatalyst was prepared by a method similar to that described in Example17 above and comprised platinum (5% by weight) and iron (0.5% by weight)on a particulate carbon support. The catalyst had not been previouslyused.

The reactor was sealed to prevent any liquid inlet or outlet flow. Thereaction mixture was heated to about 100° C. and brought to a pressureof about 100 psig with nitrogen. Oxygen flow (300 sccm) was initiatedand the reaction was run with no liquid inlet or outlet flow for about22 minutes. After the initial 22 minutes, slurry feed (70.4 g/min) wasinitiated, and reaction liquid was continuously withdrawn to maintain aconstant reactor level as indicated by the Drexelbrook level indicatordescribed in Example 21 above. The reactor system was allowed to run forabout 4300 minutes, after which time the liquid flow rate was doubled toeffectively reduce the reactor liquid residence time by a factor of two.The liquid product was analyzed with HPLC. Analytical results for thecontinuous oxidation are shown in Table 21 below. Profiles forglyphosate produced and NPMIDA reactant remaining in the product liquidare shown in FIG. 26.

TABLE 21 Oxidation Results for Example 23 Time NPMIDA Glyphosate (min)¹(wt %) (wt %) 66 0.72 3.59 130 0.76 3.61 488 0.85 3.67 994 0.89 3.681343 0.73 3.65 1476 0.76 3.63 1918 0.89 3.61 2458 0.81 3.59 2679 0.813.65 2807 0.80 3.63 3072 0.98 3.67 3893 0.88 3.62 4113 0.89 3.54 42150.86 3.56 4314 1.99 2.73 4334 2.11 2.82 ¹Time after slurry feed started.

Example 24 Continuous Oxidation of NPMIDA to Glyphosate in Two StirredTank Reactors in Series

This example demonstrates the continuous oxidation of NPMIDA toglyphosate in a continuous reactor system comprising two stirred tankreactors staged in series.

Referring to FIG. 27, the experiment was conducted in a continuousreactor system comprising two reactors and a crystallizer. The tworeactors (each 1 gallon stainless steel autoclaves from AutoclaveEngineers, Inc., Pittsburgh, Pa.) were operated continuously as stirredtank reactors in series. The continuous reactor system was arranged sothat aqueous slurry feed material was continuously introduced into thefirst reaction zone (reactor R1). Liquid effluent was continuouslywithdrawn from R1 and introduced into the second reaction zone (reactorR2). Liquid effluent was continuously withdrawn from R2 and introducedinto the crystallizer for product recovery of a glyphosate slurry.Oxygen was fed independently to each reaction zone, while product gaswas vented from each reactor independently. Oxygen gas was introducedinto R1 through a frit located near an agitator impeller (2″ turbineblade impeller). Oxygen gas was introduced into R2 in the headspaceabove the liquid level and a DISPERSIMAX type 2.5″ impeller was utilizedto effectively back-mix the headspace gas into the reaction zone. Thetemperature of the reaction mass in each reactor was controlled by aninternal cooling coil. Liquid effluent was removed from R1 via a frit,which allowed the heterogeneous catalyst to remain in R1. Similarly,liquid effluent was removed from R2 via a frit to maintain theheterogeneous catalyst inside of R2. The reaction mass/volume in eachreactor was maintained constant.

The continuous reactor system was started up in a manner similar to thatdescribed in Example 21 above in that the reactors were started in batchmode with liquid flow through the system initiated shortly afterward.The feed material was an aqueous slurry containing NPMIDA (about 7.6% byweight), glyphosate (about 2.8% by weight), formaldehyde (about 2200 ppmby weight) and formic acid (about 4500 ppm by weight). A low level ofNaCl (about 450 ppm) was also added to the feed to mimic chlorideimpurity typically present in commercially available NPMIDA. Thecatalyst was prepared by a method similar to that described in Example17 above and comprised platinum (5% by weight) and iron (0.5% by weight)on a particulate carbon support. Aqueous slurry feed material andcatalyst were charged to each reactor to give about 2% by weightcatalyst concentration in each reactor, where the target total reactormasses were 2693 grams and 1539 grams respectively for R1 and R2.

The operating conditions are summarized in Table 22 below. Analyticalresults for the aqueous slurry feed composition, R1 liquid and gaseffluent, and R2 liquid and gas effluent, which were analyzed by HPLC,are shown in Table 23 below.

TABLE 22 Summary of Operating Conditions for Example 24. R1 R2 CatalystConcentration in reactor: 2 wt % 2 wt % Agitator RPM: 1000 1200 LiquidFlow: 128 ml/min 128 ml/min Pressure: 116 psig 90 psig Oxygen Flow Rate:~1840 sccm ~390 sccm Temperature: 100° C. 105° C. Reaction Mass: 2693 g1539 g Impeller Type: radial (2″) DISPERSIMAX (2.5″)

TABLE 23 Oxidation Results for Example 24. Elapsed Time NPMIDAGlyphosate Formaldehyde Formic Acid (hrs) (wt %) (wt %) (ppm) (ppm)Reactor Feed 8.13 2.98 2348.5 5562.6 1.3 7.50 2.84 2290.0 4620.9 2.57.45 2.74 2244.2 4515.9 3.6 7.45 2.74 2244.2 4515.9 4.5 7.45 2.74 2244.24515.9 5.5 7.79 2.84 2271.0 4590.0 6.5 7.79 2.84 2271.0 4590.0 7.5 7.572.81 2286.8 4584.9 8.8 7.57 2.81 2286.8 4584.9 First Reactor (R1) Outlet1.3 0.38 7.15 385.4 6115.1 2.5 0.41 6.65 328.1 4297.7 3.6 1.18 6.83300.2 4841.8 4.5 0.79 6.56 307.2 4746.3 5.5 1.07 6.81 317.1 5193.0 6.50.88 6.48 323.6 5045.8 7.5 0.90 6.50 315.6 4976.0 8.8 1.38 6.42 323.05305.2 Second Reactor (R2) Outlet 1.3 0.03 6.84 475.9 3680.4 2.5 0.006.96 194.7 1048.1 3.6 0.02 7.23 424.0 3702.4 4.5 0.00 6.97 534.4 3006.45.5 0.00 7.27 1025.5 6176.5 6.5 0.00 6.89 1524.2 5471.0 7.5 0.01 6.971663.9 5468.1 8.8 0.03 7.07 1883.0 5808.2

Example 25 Continuous Oxidation of NPMIDA to Glyphosate in Two StirredTank Reactors in Series

This example demonstrates the continuous oxidation of NPMIDA toglyphosate in a continuous reactor system comprising two stirred tankreactors in series where the liquid effluent from the second reactor issent to a crystallizer for glyphosate recovery and the resulting motherliquor is recycled from the crystallizer back to the first reactor aspart of the reactor feed.

Referring to FIG. 28, Example 25 was conducted in a continuous reactorsystem similar to that described in Example 24 above, except that motherliquor from the crystallizer was recycled back to the first reactor R1.The continuous reactor system was started up in a manner similar to thatdescribed in Example 21 above in that the reactors were started in batchmode with liquid flow through the system initiated shortly afterward.The crystallizer (30 L) was initially charged with an aqueous slurryfeed material comprising NPMIDA (0.16% by weight), glyphosate (2.0% byweight), formaldehyde (2754 ppm by weight) and formic acid (5637 ppm byweight) and was operated at about 60° C. and 1 atm pressure. The slurryfeed system was charged with an aqueous slurry feed material comprisingNPMIDA (about 25% by weight). The catalyst used was similar to theheterogeneous particulate catalyst used in Example 24.

Aqueous slurry and catalyst were charged to each reactor to give about2% by weight catalyst concentration in each reactor, where the targettotal reactor masses were 2693 grams and 1539 grams respectively for R1and R2. After the initial batch runs, liquid flow through the system wasinitiated. Liquid entering R1 comprised the aqueous slurry feed material(about 40 ml/min) and mother liquor recycle (about 80 ml/min) from thecrystallizer. The liquid level in each reactor was controlled during therun to maintain a constant reaction mass in each reactor, targetinghydraulic residence times of 21 minutes and 12.2 minutes in R1 and R2,respectively, and giving a total liquid flow through the system of about120 ml/min.

The operating conditions are summarized in Table 24 below. The liquidproduct was analyzed with HPLC. Analytical results for the aqueousslurry feed composition, R1 liquid effluent, and R2 liquid effluent areshown in Table 25 below. The elapsed time refers to the time periodafter which continuous liquid flow was initiated.

TABLE 24 Summary of Operating Conditions for Example 25. R1 R2 CatalystConcentration in reactor: 2 wt % 2 wt % Agitator RPM: 1015 1005 CombinedLiquid Flow 121 ml/min 121 ml/min Through System: Pressure: 116 psig 89psig Oxygen Flow Rate: ~1660 sccm ~280 sccm Temperature: 100° C. 106° C.Reaction Mass: 2545 g 1592 g Impeller Type: radial (2″) DISPERSIMAX 2.5″

TABLE 25 Oxidation Results for Example 25. Elapsed Time NPMIDAGlyphosate Formaldehyde Formic Acid (hrs) (wt %) (wt %) (ppm) (ppm)First Reactor (R1) Outlet 1.3 1.50 6.99 849.8 3202.0 2.9 0.45 8.161053.5 2789.3 4.1 0.62 8.40 1199.4 3178.0 5.0 0.65 8.07 1240.8 3348.66.1 1.21 7.51 1294.7 3701.1 Second Reactor (R2) Outlet 1.3 2.11 6.50374.2 1682.3 2.9 0.27 8.02 501.0 2171.4 4.1 0.15 8.55 451.0 2678.0 5.00.12 8.49 564.4 3107.5 6.1 0.19 8.02 577.3 3505.7

Example 26 Continuous Oxidation of NPMIDA to Glyphosate in the Presenceof a Pt/Sn/C Catalyst

This example demonstrates the continuous oxidation of NPMIDA toglyphosate in the presence of a Pt/Sn/C heterogeneous particulatecatalyst in a stirred tank reactor. The experiment was designed tosimulate conditions that might prevail in a second reaction zone of acontinuous reactor system.

The experiment was conducted in a continuous reactor system comprising a500 ml Hastelloy C autoclave (Autoclave Engineers, Inc., Pittsburgh,Pa.). The reactor was equipped with an agitator having a 1.25″ diameterradial six-blade turbine impeller. The liquid level in the reactor wasmonitored using a level indicator similar to that described in Example21 above. An internal cooling coil was utilized to control thetemperature within the reactor during the course of the reaction.

During operation, the reactor was continuously fed a gaseous stream ofoxygen and an aqueous slurry feed material containing NPMIDA. The oxygenwas introduced into the reaction medium through a frit located near theimpeller. A liquid product stream containing glyphosate product wascontinuously withdrawn from the reactor through a frit, which allowedany catalyst charged to the reactor to remain in the reaction medium.The withdrawn liquid product stream was mixed in-line with a basicsolution capable of dissolving glyphosate. The product gas stream(containing CO₂ and unreacted oxygen) was continuously vented from thereactor headspace.

The continuous reactor system was started up in a manner similar to thatdescribed in Example 21 above in that the reactor was started in batchmode with liquid flow through the system initiated shortly afterward.The aqueous slurry feed material comprised NPMIDA (2.46% by weight),glyphosate (3.72% by weight), formaldehyde (1381 ppm by weight) andformic acid (6485 ppm by weight). The catalyst was prepared by a methodsimilar to that described in Example 14 above and comprised platinum (5%by weight) and tin (1.0% by weight) on a particulate carbon support.

The operating conditions are summarized in Table 26 below. The liquidproduct was analyzed with HPLC. Analytical data from the oxidation runare shown in Table 27 below.

TABLE 26 Summary of Operating Conditions for Example 26. CatalystConcentration in Reaction Mass: 1 wt % Agitator RPM: 1000 Liquid Flow:30.8 ml/min Pressure: 100 psig Gas Flow Rate: 270 sccm Temperature: 100°C. Reaction Mass: 300 g

TABLE 27 Oxidation Results for Example 26. Feed Composition — NPMIDA (wt%) Glyphosate (wt %) — 2.46 3.72 Reactor Effluent Elapsed Time (mins)NPMIDA (wt %) Glyphosate (wt %)  120 0.07 5.47 1200 0.09 5.58 2500 0.125.45 3500 0.15 5.47

Example 27 Continuous Oxidation of NPMIDA to Glyphosate in the Presenceof a Pt/Sn/C Catalyst

This example demonstrates the continuous oxidation of NPMIDA toglyphosate in the presence of a Pt/Sn/C heterogeneous particulatecatalyst in a stirred tank reactor. The experiment was designed tosimulate conditions that might prevail in a second reaction zone of acontinuous reactor system. Also, the oxygen flow rate was varied toillustrate the impact of various oxygen flow rates on conversion.

The experiment was conducted in a continuous reactor system similar tothe reactor system described in Example 26 above. During operation, thereactor was continuously fed a gaseous stream of oxygen and an aqueousslurry feed material containing NPMIDA. The oxygen was introduced intothe reaction medium through a frit located near the impeller. A liquidproduct stream containing glyphosate product was continuously withdrawnfrom the reactor through a frit, which allowed any catalyst charged tothe reactor to remain in the reaction medium. The withdrawn liquidproduct stream was mixed in-line with a basic solution capable ofdissolving glyphosate. The product gas stream (containing CO₂ andunreacted oxygen) was continuously vented from the reactor headspace.

The continuous reactor system was started up in a manner similar to thatdescribed in Example 21 above in that the reactor was started in batchmode with liquid flow through the system initiated shortly afterward.The aqueous slurry feed material comprised NPMIDA (about 2.8% byweight), glyphosate (about 4.2% by weight), formaldehyde (about 1425 ppmby weight) and formic acid (about 6570 ppm by weight). The catalyst wasprepared by a method similar to that described in Example 14 above andcomprised platinum (5% by weight) and tin (1.0% by weight) on aparticulate carbon support.

The operating conditions are summarized in Table 28 below. During thecourse of this experiment, the oxygen flow rate to the reactor wasramped up and ramped back down over the range of from 75 to 300 sccm.The liquid product was analyzed with HPLC. Analytical data from thecontinuous oxidation are shown in Table 29 below.

TABLE 28 Operating Conditions for Example 27 Catalyst Concentration inreactor: 1 wt % Agitator RPM: 1000 Liquid Flow: 30 ml/min Pressure: 10psig Oxygen Flow Rate: variable (75-300 sccm) Temperature: 100° C.Reaction Mass: 300 g Impeller Type: radial (1.25″)

TABLE 29 Oxidation Results for Example 27 Feed Composition NPMIDAGlyphosate Formaldehyde Formic (wt %) (wt %) (ppm) Acid (ppm) 2.83 4.171425.1 6569.8 Reactor Outlet Composition Formic O2 Elapsed NPMIDAGlyphosate Formaldehyde Acid Flow Time (hrs) (wt %) (wt %) (ppm) (ppm)(sccm) 0.0 1.1 4.59 1699.5 5463.1 74.7 0.4 1.8 4.86 1543.7 6067.0 49.70.7 1.98 4.74 1431.7 6020.5 49.8 1.0 2.02 4.90 1478.1 6105.2 52.7 1.41.97 4.80 1474.0 6209.0 54.7 1.7 1.91 4.73 1441.3 5806.0 54.7 2.0 1.674.93 1588.8 6006.9 74.7 2.4 1.54 5.03 1590.2 6135.3 74.7 2.7 1.63 5.201625.7 6280.1 74.7 3.0 1.64 5.19 1591.5 6015.1 74.8 3.4 1.61 5.00 1547.85834.7 74.7 3.7 1.61 5.12 1541.0 5864.8 74.7 4.0 1.58 5.15 1566.9 5791.074.7 4.4 1.61 5.23 1565.6 6274.6 74.7 4.7 0.66 6.01 2099.7 6337.5 149.85.0 0.51 6.20 2109.3 6036.9 149.6 5.4 0.46 5.81 1976.8 5688.5 149.8 5.70.47 6.04 2094.3 5849.7 149.8 6.0 0.45 6.04 2109.3 5785.5 149.8 6.4 0.456.15 2157.1 6101.1 149.8 6.7 0.41 5.70 2016.4 5489.1 149.8 7.0 0.38 5.381907.1 5213.1 149.8 7.4 0.41 5.79 2056.0 5531.4 149.8 7.7 0.44 6.262230.9 5949.5 149.8 8.0 0.35 6.43 2337.4 6083.4 166.0 8.4 0.48 6.092356.6 6147.6 210.6 8.7 0.33 6.37 2665.3 6464.5 224.9 9.0 0.34 6.242684.4 6308.8 224.9 9.4 0.36 6.30 2741.8 6412.6 224.9 9.7 0.19 6.582680.3 6340.2 224.7 10.0 0.22 6.54 2530.1 6367.5 224.7 10.4 0.20 6.522560.1 6256.9 224.7 10.7 0.18 5.51 2163.9 5241.8 224.7 11.0 0.22 6.372502.7 6202.2 224.7 11.4 0.23 6.73 2648.9 6449.5 224.7 11.7 0.20 6.352517.8 6131.2 224.7 12.0 0.16 5.11 1987.7 4889.4 224.7 12.4 0.20 6.042430.3 5877.1 224.7 12.7 0.13 6.67 2777.3 6276.0 299.7 13.0 0.13 6.732844.3 6349.8 299.7 13.4 0.15 6.61 2808.8 6204.9 299.7 13.7 0.08 5.572323.8 5144.8 299.7 14.0 0.10 6.61 2704.9 6215.9 299.8 14.4 0.14 6.802774.6 5810.1 299.8 14.7 0.13 6.89 2845.6 6147.6 299.8 15.0 0.12 6.862871.6 6232.3 299.8 15.3 0.11 6.53 2745.9 5874.3 299.8 15.7 0.13 6.282668.0 5654.4 299.8 16.0 0.15 6.86 2923.5 6360.7 299.8 16.3 0.16 6.862970.0 6702.2 299.8 16.7 0.17 6.57 2874.3 6459.0 224.8 17.0 0.23 6.532789.6 6508.2 224.8 17.3 0.24 6.57 2822.4 6403.0 224.8 17.7 0.25 6.632822.4 6580.2 224.8 18.0 0.23 6.39 2736.3 6385.3 224.8 18.3 0.22 6.192668.0 6189.9 224.8 18.7 0.23 6.53 2811.5 6546.5 224.8 19.0 0.24 6.522792.4 6445.4 224.8 19.3 0.24 6.20 2655.7 6138.0 224.8 19.7 0.35 6.492752.7 6278.7 224.8 20.0 0.48 6.23 2572.4 6460.4 149.8 20.3 0.53 6.152513.7 6030.1 149.8 20.7 0.50 6.34 2542.4 6143.5 149.8 21.0 0.51 6.312527.3 6113.4 149.8 21.3 0.48 6.31 2527.3 6050.6 149.8 21.7 0.48 6.422523.2 5885.3 149.8 22.0 0.46 6.16 2430.3 5655.8 149.8 22.3 0.48 6.382521.9 6032.8 149.8 22.7 0.45 6.12 2426.2 5695.4 149.8 23.0 0.46 6.262480.9 5868.9 149.8 23.3 1.18 6.20 2117.5 6220.0 74.8 23.7 1.30 5.871956.3 5970.0 74.8 24.0 1.61 5.68 1916.7 5909.9 74.8 24.3 1.50 5.611795.1 5720.0 74.8 24.7 1.61 5.85 1847.0 5862.0 74.8 25.0 1.68 5.871908.5 6599.8 74.8 25.3 1.69 5.83 1868.9 6653.0 74.8 25.7 1.60 5.571773.2 6460.4 74.8 26.0 1.71 5.75 1837.4 6577.9 74.8 26.3 1.60 5.461751.4 6299.2 74.8 26.7 1.65 5.71 1827.9 6416.7 74.8 27.0 1.64 5.601811.5 6433.1 74.8 27.3 1.63 5.63 1826.5 6297.8 74.8

Example 28 Continuous Oxidation of NPMIDA to Glyphosate in Two StirredTank Reactors in Series

This example demonstrates the continuous oxidation of NPMIDA toglyphosate in a continuous reactor system comprising two stirred tankreactors in series. In this example, the particulate heterogenouscatalyst was transferred from the first reactor to the second reactor.The catalyst exited the second reactor with the second reactor liquideffluent, was separated by filtration, and recycled back to the firstreactor zone.

Referring to FIG. 29, the reaction was performed in a continuous reactorsystem comprising two stirred tank reactors, a slurry feed system and acatalyst filtration system. The two reactors (each 1 gallon stainlesssteel autoclaves from Autoclave Engineers, Inc., Pittsburgh, Pa.) wereoperated continuously as stirred tank reactors in series. Oxygen was fedto each respective reactor. Liquid effluent was withdrawn from the firstreactor (R1) through a dip-tube which allowed for catalyst to beentrained with the liquid effluent from R1 to the second reactor (R2).Some reaction product gases were also entrained in the dip-tube from R1to R2, and other reaction product gases in R1 were vented from thereactor. Similarly, liquid effluent was withdrawn from R2 through adip-tube, which allowed for catalyst and some reaction product gases tobe removed with the effluent. R2 liquid effluent which contained thecatalyst was transferred to a catalyst filtration system. The catalystfiltration system generated an uninterrupted flow of catalyst-freefiltrate as product. A filter back wash was fed to the catalystfiltration system to wash the filtered catalyst back to R1 in acontinuous fashion. Oxygen was introduced into R1 and R2 through fritswhich were each located near an agitator impeller (2″ turbine bladeimpeller in both R1 and R2). An internal cooling coil was used tocontrol the temperature in each reactor.

An aqueous slurry feed material containing about 25% by weight NPMIDAwas fed to the reactor at a rate of about 50 ml/min. The filter backwash contained NPMIDA (about 3% by weight), glyphosate (about 0.1% byweight), formaldehyde (about 3000 ppm by weight), and formic acid (about7000 ppm by weight). The filter back wash was returned to R1 at a rateof about 100 ml/min. The catalyst was prepared by a method similar tothat described in Example 17 above and comprised platinum (5% by weight)and iron (0.5% by weight) on a particulate carbon support. Catalyst wascharged to the reactor system to provide an initial concentration ofabout 1% by weight catalyst. The reactor system was started up in amanner similar to that described in Example 24 in that the reactors werestarted in batch mode with liquid flow through the system initiatedshortly thereafter.

The operating conditions are summarized in Table 30 below. The liquidproduct was analyzed with HPLC. Oxidation results are shown in Table 31below. Table 31 gives data describing the inlet stream composition intoR1 (including the combined from the catalyst filtration back-wash andthe aqueous slurry feed) and liquid effluent compositions for R1 and R2.

TABLE 30 Summary of Operating Conditions for Example 28 R1 R2 CatalystConcentration in reactor: ~1 wt % ~1 wt % Agitator RPM: ~1000 ~1000Liquid Flow (combined through ~150 ml/min ~150 ml/min reactors):Pressure: 120-140 psig 120-140 psig Oxygen Flow Rate: 2000-2500 sccm~400 sccm Temperature: ~100-105° C. ~105° C. Reaction Mass: ~2950 g~1726 g Impeller Type: radial (2″) radial (2″)

TABLE 31 Oxidation Results for Example 28 Elapsed NPMIDA GlyphosateFormaldehyde Formic Time (hrs) (wt %) (wt %) (ppm) Acid (ppm) FeedComposition 20.9 8.08 1.81 1778.0 4650.0 23.8 8.38 1.80 1757.2 4400.530.3 8.54 1.77 1518.5 4560.5 33.8 8.55 1.78 1522.8 4573.5 First Reactor(R1) Outlet 20.9 3.50 7.87 668.0 4384.2 23.8 3.44 7.32 651.6 4425.4 30.32.87 8.43 756.7 5057.3 33.8 2.71 8.39 798.8 5206.4 Second Reactor (R2)Outlet 20.9 2.55 8.04 302.5 3382.5 23.8 3.27 7.71 258.7 3791.2 30.3 0.406.98 — 4017.3 33.8 1.81 8.24 312.5 3945.7

Example 29 Continuous Oxidation of NPMIDA to Glyphosate in the Presenceof a Pt/Sn/C Catalyst

This example demonstrates the continuous oxidation of NPMIDA toglyphosate in the presence of a Pt/Sn/C heterogeneous particulatecatalyst. The experiment was designed to simulate reaction conditionsthat might prevail in a first stirred tank reaction zone of a continuousreactor system.

The experiment was conducted in continuous reactor system similar tothat described in Example 26 with the exception that a 1000 ml HastelloyC autoclave was used. The reactor was equipped with an agitator having a1.25″ diameter radial six-blade turbine impeller. The liquid level inthe reactor was monitored using a level indicator similar to that usedin Example 21. An internal cooling coil was utilized to control thetemperature within the reactor during the course of the reaction.

During operation, the reactor was continuously fed a gaseous stream ofoxygen and an aqueous slurry feed material containing NPMIDA. The oxygenwas introduced into the reaction medium through a frit located near theimpeller. A liquid product stream containing glyphosate product wascontinuously withdrawn from the reactor through a frit, which allowedany catalyst charged to the reactor to remain in the reaction medium.The withdrawn liquid product stream was then mixed in-line with a basicsolution capable of dissolving glyphosate. The product gas stream(containing CO₂ and unreacted oxygen) was continuously vented from thereactor headspace.

The aqueous slurry feed material comprised NPMIDA (about 7.7% byweight), formaldehyde (about 3000 ppm by weight) and formic acid (about6100 ppm by weight). The catalyst was prepared by a method similar tothat described in Example 14 above and comprised platinum (5% by weight)and tin (1.0% by weight) on a particulate carbon support. The continuousreactor system was started up in a manner similar to that described inExample 21 above in that the reactor was started in batch mode withliquid flow through the system initiated shortly thereafter. Theoperating conditions are summarized in Table 32 below. Analytical datafrom the continuous oxidation are shown in Table 33 below.

TABLE 32 Summary of Operating Conditions for Example 29. CatalystConcentration in reactor: 1 wt % Agitator RPM: 1000 Liquid Flow: 30.8ml/min Pressure: 100 psig Oxygen Flow Rate: 647 sccm Temperature: 100°C. Reaction Mass: 725 g Impeller Type: radial (1.25″)

TABLE 33 Oxidation Results for Example 29. Feed Composition NPMIDA (wt%) Glyphosate (wt %) 7.73 0.00 Reactor Outlet Elapsed Time (hrs) NPMIDA(wt %) Glyphosate (wt %) 0.0 0.04 1.35 0.5 0.29 4.42 1.0 0.34 4.91 1.50.41 5.18 2.0 0.58 5.55 2.5 0.97 6.50

Example 30 Continuous Oxidation of NPMIDA to Glyphosate in the Presenceof a Pt/Sn/C Catalyst

This example demonstrates the continuous oxidation of NPMIDA toglyphosate in the presence of a Pt/Sn/C heterogeneous particulatecatalyst in a continuous reactor system having a single reaction zone.

The experiment was conducted in a continuous reactor system similar tothat described in Example 27 above comprising a 500 ml Hastelloy Cautoclave (Autoclave Engineers, Inc., Pittsburgh, Pa.). The reactor wasequipped with an agitator having a 1.25″ diameter radial six-bladeturbine impeller. The liquid level in the reactor was monitored using alevel indicator similar to that described in Example 21 above. Aninternal cooling coil was utilized to control the temperature within thereactor during the course of the reaction.

During operation, the reactor was continuously fed a gaseous stream ofoxygen and an aqueous slurry feed material containing NPMIDA. The oxygenwas introduced into the reaction medium through a frit located near theimpeller. A liquid product stream containing glyphosate product wascontinuously withdrawn from the reactor through a frit, allowing anycatalyst charged to the reactor to remain in the reaction medium. Theproduct gas stream (containing CO₂ and unreacted oxygen) wascontinuously vented from the reactor headspace.

The continuous reactor system was started up in a manner similar to thatdescribed in Example 21 above in that the reactor was started in batchmode with liquid flow through the system initiated shortly afterward.The aqueous slurry feed material comprised NPMIDA (about 2.9 wt %). Thecatalyst was prepared by a method similar to that described in Example14 above and comprised platinum (5% by weight) and tin (1.0% by weight)on a particulate carbon support. The operating conditions are summarizedbelow in Table 34. The liquid product was analyzed with HPLC. Analyticaldata from the continuous oxidation are shown in Table 35 below.

TABLE 34 Summary of Operating Conditions for Example 30 CatalystConcentration in reactor: 1 wt % Agitator RPM: 1000 Liquid Flow: 15.3ml/min Pressure: 100 psig Oxygen Flow Rate: 150 sccm Temperature: 95° C.Reaction Mass: 300 g Impeller Type: radial (1.25″)

TABLE 35 Oxidation Results for Example 30 Elapsed Time NPMIDA GlyphosateFormaldehyde Formic Acid (hrs) (wt %) (wt %) (ppm) (ppm) 0.0 2.87 0.004.6 14.9 0.3 2.94 0.01 13.4 18.9 0.7 2.01 0.79 760.4 563.1 1.0 0.12 2.071893.6 1566.6 1.3 0.07 2.32 1953.6 1713.4 1.7 0.01 2.27 2111.1 1497.22.0 0.00 2.27 2167.1 1487.9 2.3 0.00 2.26 2155.1 1509.2 2.7 0.00 2.262183.1 1495.9 3.0 0.00 2.27 2189.8 1549.3 3.3 0.00 2.27 2195.1 1535.93.7 0.00 2.28 2196.5 1538.6 4.0 0.04 2.26 2184.5 1522.6 4.3 0.03 2.262184.5 1474.5 4.4 0.00 2.26 2177.8 1478.5

Example 31 Continuous Oxidation of NPMIDA to Glyphosate in the Presenceof a Pt/Sn/C Catalyst

This example demonstrates the continuous oxidation of NPMIDA toglyphosate in the presence of a Pt/Sn/C heterogeneous particulatecatalyst. The experiment was designed to illustrate the impact ofpressure, liquid and gas flow on conversion of NPMIDA to glyphosate.

The experiment was conducted in a continuous reactor system comprising a500 ml Hastelloy C autoclave (Autoclave Engineers Inc.) similar to thatdescribed in Example 30. The reactor was equipped with an agitatorhaving a 1.25″ diameter radial six-blade turbine impeller. The liquidlevel in the reactor was monitored using a level indicator similar tothat used in Example 21. An internal cooling coil was utilized tocontrol the temperature within the reactor during the course of thereaction.

During operation, the reactor was continuously fed a gaseous stream ofoxygen and an aqueous slurry feed material containing NPMIDA. The oxygenwas introduced into the reaction medium through a frit located near theimpeller. A liquid product stream containing glyphosate product wascontinuously withdrawn from the reactor through a frit, allowing anycatalyst charged to the reactor to remain in the reaction medium. Theproduct gas stream (containing CO₂ and unreacted oxygen) wascontinuously vented from the reactor headspace.

The aqueous slurry feed material comprised NPMIDA (about 3.0% byweight), formaldehyde (about 1000 ppm by weight) and formic acid (about5100 ppm by weight). The catalyst was prepared by a method similar tothat described in Example 14 above and comprised platinum (5% by weight)and tin (1.0% by weight) on a particulate carbon support.

The continuous reactor system was started up in a manner similar to thatdescribed in Example 21 above in that the reactor was started in batchmode with liquid flow through the system initiated shortly thereafter.During the course of the experiment, the oxygen flow rate to the reactorwas ramped up and ramped back down over the range of 75 to 300 sccm. Theoperating conditions are summarized below in Table 36. The liquidproduct was analyzed with HPLC. Analytical data from the continuousoxidation are shown in Table 37 below.

TABLE 36 Summary of Operating Conditions for Example 31 CatalystConcentration in reactor: 1 wt % Agitator RPM: 1000 Liquid Flow:variable (see data) Pressure: variable (see data) Oxygen Flow Rate:variable (see data) Temperature: 100° C. Reaction Mass: 300 g ImpellerType: radial (1.25″)

TABLE 37 Oxidation Results for Example 31 Liquid O₂ Feed Flow % CO₂ TimeRates Pressure Rate GI Gly HCHO HCOOH NMG NFG AMPA MAMPA in Off- (min)(ml/min) (psig) (sccm) (wt %) (wt %) (ppm) (ppm) (ppm) (ppm) (ppm) (ppm)gas 0 50 100.2 149.6 0.69 1.40 1713 4901 26 167 20 11 66 21 50 100.1149.6 0.88 1.35 1756 5247 31 139 12 12 74 42 50 99.9 149.6 0.92 1.341709 5342 19 136 9 8 76 63 50 99.9 149.6 0.90 1.34 1671 5293 23 133 9 1278 84 50 100.1 149.6 0.91 1.40 1709 5411 23 135 9 13 79 105 50 100 149.60.81 1.41 1685 5311 21 128 12 12 82 126 50 99.9 149.6 0.84 1.42 16865416 24 128 11 12 84 147 50 100 149.6 0.85 1.42 1685 5365 21 126 9 12 84168 50 100 149.6 0.88 1.41 1650 5357 20 120 9 12 85 189 50 100 149.60.87 1.42 1651 5344 19 117 9 9 85 210 50 100.1 149.6 0.94 1.36 1607 531220 115 8 9 83 231 35 100 149.6 0.38 1.56 1312 4661 21 136 15 12 87 25235 100.6 149.6 0.45 1.84 1799 5456 24 173 18 11 83 273 35 100 149.6 0.331.93 1794 5456 33 213 29 19 80 294 35 99.9 149.6 0.32 1.88 1785 5335 13199 27 — 85 315 35 100.1 149.6 0.31 1.90 1834 5388 30 205 27 13 85 33635 100.2 149.6 0.30 1.90 1866 5398 32 184 26 19 84 357 35 100 149.6 0.361.92 1887 5421 35 182 26 19 84 378 35 100 149.6 0.33 1.92 1884 5322 31175 27 17 83 399 35 100 149.6 0.36 1.92 1906 5366 33 176 26 18 83 420 3599.8 149.6 0.35 1.95 1932 5386 33 177 27 19 83 441 35 100 149.6 0.341.93 1950 5331 33 174 29 18 83 462 35 100.1 149.6 0.33 1.90 1956 5248 33170 31 18 83 483 20 100.2 149.6 0.04 2.18 1547 4205 9 233 188 −52 77 50420 100.2 149.6 0.01 2.15 1656 4325 10 231 304 75 68 525 20 100 149.60.01 2.19 1756 4613 12 239 329 80 66 546 20 99.9 149.6 0.01 2.15 17994703 12 233 324 84 64 567 20 100 149.6 0.01 2.13 1822 4743 15 226 319 7963 588 20 100 149.6 0.02 2.10 1839 4750 16 229 309 80 62 614 — 99.9149.6 — — — — — — 61 635 — 100.1 149.6 — — — — — 61 656 — 99.9 149.6 — —— — — — 60 677 50 100 149.6 0.52 1.79 2379 5668 27 181 66 23 64 698 3599.8 149.6 0.66 1.14 1563 4079 20 88 12 13 70 719 35 100 149.6 0.29 1.952271 5382 34 180 38 18 80 740 35 99.5 149.6 0.27 2.00 2334 5428 35 18339 19 79 761 35 100.8 0 — — — — — — 54 782 35 100.1 0 2.76 0.10 980 493012 13 0 8 53 803 35 99.7 150 1.67 1.07 1344 4074 8 97 5 9 17 824 35100.2 150 0.65 1.71 2002 5269 20 148 17 12 59 845 35 100.2 150 0.55 1.792218 5499 24 151 23 14 68 979 35 99.6 199.8 0.02 1.93 2271 4947 20 174832 32 44 1047 35 100 199.8 0.38 1.90 2342 5881 29 152 25 20 43 1067 35100 199.8 0.19 2.05 2532 5723 34 162 50 22 56 1086 35 99.9 199.8 0.172.02 2663 5615 35 164 57 25 57 1106 35 100 199.8 0.19 2.02 2703 5584 36163 55 25 56 1126 35 100.2 199.8 0.19 2.05 2797 5672 37 165 54 26 551145 35 100.1 199.8 0.22 1.99 2783 5551 37 158 48 24 54 1165 35 100.1199.8 0.23 1.99 2813 5567 36 161 45 26 54 1186 35 100.1 199.8 0.24 2.012835 5605 35 158 44 25 53 1205 35 100 199.8 0.25 2.01 2880 5592 36 16245 25 53 1225 35 100.2 199.8 0.25 2.00 2839 5535 35 158 43 24 52 1244 35100.1 199.8 0.25 1.96 2809 5478 35 156 43 23 53 1274 20 100.1 199.8 0.002.05 1740 4227 8 197 444 92 52 1296 20 100 199.8 0.00 2.07 1955 4675 16208 439 86 48 1317 20 100 199.8 0.00 2.05 1872 4701 12 172 439 92 481337 20 99.8 199.8 0.00 2.05 1688 4991 15 213 402 87 47 1354 20 100.2199.8 0.00 2.07 2005 4982 14 217 417 86 46 1374 20 100.3 199.8 0.00 2.102122 5160 19 219 418 72 46 1444 20 99.9 199.8 0.00 2.09 2183 5228 21 220369 93 44 1464 20 99.5 199.8 0.00 2.10 2127 5183 17 224 407 89 44 148435 100 199.8 0.22 2.05 2966 5982 50 184 69 30 47 1504 35 100.1 199.80.28 1.95 2930 5759 37 163 38 28 49 1524 35 99.8 199.8 0.34 1.93 30415728 48 158 38 22 47 1544 35 100 199.8 0.36 1.99 3003 5685 30 150 38 2247 1564 35 99.4 199.8 0.15 2.15 2965 5668 34 162 63 27 57 1584 35 100199.8 0.26 2.00 3027 5687 45 167 46 27 52 1604 35 100.2 199.8 0.26 2.003020 5647 39 165 44 26 50 1624 35 129.7 199.8 0.15 2.07 3116 5613 42 16868 30 51 1644 35 130 199.8 0.14 2.07 3146 5571 41 169 70 30 51 1664 35130.1 199.8 0.16 2.06 3166 5599 41 168 63 30 50 1684 35 129.9 199.8 0.152.07 3199 5626 41 170 64 30 50 1704 35 129.8 199.8 0.16 2.07 3179 564041 168 58 30 49 1724 35 130 199.8 0.17 2.06 3189 5674 40 165 58 27 481744 35 129.8 199.8 0.19 2.02 3178 5669 41 164 54 28 48 1764 35 129.8199.8 0.20 2.00 3179 5625 41 163 52 26 47 1784 35 130 199.8 0.30 1.992839 5551 31 161 55 22 52 1804 35 130.2 199.8 0.20 2.01 2657 5350 30 15968 31 56 1824 35 130 199.8 0.10 2.14 2844 5516 31 177 100 38 56 1844 35130 199.8 0.10 2.13 2822 5474 30 337 99 37 57 1864 35 130 150 0.20 2.052711 5517 35 313 64 29 62 1884 35 129.9 150 0.23 2.08 2754 5580 30 31442 23 65 1904 35 130.4 150 0.36 1.90 2561 5336 32 274 36 20 66 1924 35129.4 150 0.43 1.89 2691 5498 31 275 33 18 64 1944 35 130 150 0.22 2.092583 5419 40 314 77 31 73 1964 35 130 150 0.28 1.99 2789 5505 35 294 4222 68 1984 35 130.2 150 0.27 1.98 2885 5510 35 292 43 23 65 2004 35129.9 150 0.27 1.99 2926 5529 34 295 42 25 64 2024 35 130.1 150 0.242.02 2960 5525 37 301 44 27 63 2044 35 130.2 150 0.25 2.04 2989 5607 35313 44 25 62 2064 35 129.7 150 0.30 1.98 2995 5549 34 299 40 21 62 208435 129.9 150 0.28 1.99 3033 5588 35 305 40 23 61 2104 20 130.4 200 0.182.10 2637 5240 25 307 229 62 54 2124 20 130 200 0.00 2.12 2078 4637 13283 634 93 48 2144 20 129.7 200 0.00 2.20 2171 4975 14 332 686 114 462164 20 129.8 200 0.00 2.12 2299 5001 24 219 606 100 45 2184 20 129.7200 0.00 2.13 2237 4955 20 230 692 110 46 2204 20 130.1 200 0.00 2.142260 5044 19 238 695 123 46 2224 20 130.4 200 0.00 2.12 2319 5014 23 230662 111 45 2244 20 129.7 200 0.00 2.14 2386 5142 25 235 644 116 44 226420 129.9 200 0.00 2.12 2525 5172 34 231 516 124 42 2284 20 130 150 0.002.19 2428 5249 24 244 595 116 49 2304 20 130.3 150 0.01 2.10 2396 505831 225 460 89 53 2324 20 129.8 150 0.03 1.89 2401 4766 40 182 285 72 522344 20 129.7 150 0.01 2.15 2565 5195 37 231 377 84 51 2364 20 129.9 1500.01 2.13 2490 5064 34 225 403 89 54 2384 20 129.9 150 0.01 2.10 23824983 30 224 444 88 55 2404 20 129.7 150 0.01 2.17 2409 5063 26 228 45084 54 2424 20 129.7 150 0.01 2.18 2501 5178 30 234 418 83 54 2444 20130.3 150 0.01 2.14 2668 5358 47 231 328 75 52

Example 32 Continuous Oxidation of NPMIDA to Glyphosate in the Presenceof a Pt/Fe/C Catalyst

This example demonstrates the continuous oxidation of NPMIDA toglyphosate in the presence of a Pt/Fe/C heterogeneous particulatecatalyst in a stirred tank reactor. The experiment was designed tosimulate conditions that might prevail in a first reaction zone of acontinuous oxidation reactor system.

The experiment was conducted in a continuous reactor system similar tothat used in Example 29 where the reactor comprised a 1000 ml HastelloyC autoclave. The reactor was equipped with an agitator having a 1.25″diameter radial six-blade turbine impeller. The liquid level in thereactor was monitored using a level indicator similar to that used inExample 21. An internal cooling coil was utilized to control thetemperature within the reactor during the course of the reaction.

During operation, the reactor was continuously fed an aqueous slurryfeed material comprising NPMIDA and a gaseous stream of oxygen. Theoxygen was introduced into the reaction medium through a frit locatednear the impeller. A liquid product stream containing glyphosate productwas continuously withdrawn from the reactor through a frit, whichallowed any catalyst charged to the reactor to remain in the reactionmedium. The product gas stream (containing CO₂ and unreacted oxygen) wascontinuously vented from the reactor headspace.

The aqueous slurry feed material fed to the reactor comprised NPMIDA(9.9% by weight), glyphosate (1.3% by weight), formaldehyde (3600 ppm byweight) and formic acid (6200 ppm by weight). The catalyst was preparedby a method similar to that described in Example 17 above and comprisedplatinum (5% by weight) and iron (0.5% by weight) on a particulatecarbon support. The continuous reactor system was started up in a mannersimilar to that described in Example 21 above in that the reactor wasstarted in batch mode with liquid flow through the system initiatedshortly afterward. The operating conditions are summarized in Table 38below. The liquid product was analyzed with HPLC. Analytical data fromthe continuous oxidation are shown in Table 39 below.

TABLE 38 Summary of Operating Conditions for Example 32 CatalystConcentration in reactor: 1 wt % Agitator RPM: 1000 Liquid Flow: 35ml/min Pressure: 100 psig Oxygen Flow Rate: 630 sccm Temperature: 100°C. Reaction Mass: 725 g Impeller Type: radial (1.25″)

TABLE 39 Oxidation Results for Example 32 NPMIDA Glyphosate Time (min)(wt %) (wt %) 60 0.12 5.16 90 0.05 6.34 151 0.03 3.64 181 0.32 6.06 2110.35 6.33 241 0.34 6.23 271 0.27 6.06 301 0.32 6.22 331 0.31 6.22 3620.28 6.25 392 0.29 6.08 422 0.30 6.22 452 0.25 6.17 482 0.03 5.59 5120.01 4.03 542 0.04 4.42 573 0.18 4.84 603 0.15 5.69 633 0.23 5.85 6630.32 6.37

Example 33 Continuous Oxidation of NPMIDA to Glyphosate in the Presenceof a Pt/Fe/C Catalyst

This example demonstrates the continuous oxidation of NPMIDA toglyphosate in the presence of a Pt/Fe/C heterogeneous particulatecatalyst in a stirred tank reactor. The experiment was designed tosimulate conditions that might prevail in a second reaction zone of acontinuous reactor system.

The experiment was conducted in a continuous reactor system similar tothat used in Example 29 where the reactor comprised a 1000 ml HastelloyC autoclave. The reactor was equipped with an agitator having a 1.25″diameter radial six-blade turbine impeller. The liquid level in thereactor was monitored using a level indicator similar to that used inExample 21. An internal cooling coil was utilized to control thetemperature within the reactor during the course of the reaction.

During operation, an aqueous slurry feed material comprising NPMIDA anda gaseous stream of oxygen were continuously introduced to the reactorsystem. The oxygen was introduced into the reaction medium through afrit located near the impeller. A liquid product stream comprisingglyphosate product was continuously withdrawn from the reactor through afrit, which allowed any catalyst charged to the reactor to remain in thereaction medium. The product gas stream (containing CO₂ and unreactedoxygen) was continuously vented from the reactor headspace.

The aqueous slurry feed material contained NPMIDA (1.9% by weight),glyphosate (6.7% by weight), formaldehyde (2400 ppm by weight), formicacid (4600 ppm by weight), NMG (280 ppm by weight), AMPA (400 ppm byweight) and MAMPA (200 ppm by weight). The catalyst was prepared by amethod similar to that described in Example 17 above and comprisedplatinum (5% by weight) and iron (0.5% by weight) on a particulatecarbon support. The continuous reactor system was started up in a mannersimilar to that described in Example 21 above in that the reactor wasstarted in batch mode before initiating liquid flow through the system,which commenced shortly thereafter. The operating conditions aresummarized in Table 40 below. The liquid product was analyzed with HPLCand analytical data from the continuous oxidation are shown in Table 41.

TABLE 40 Summary of Operating Conditions for Example 33 CatalystConcentration in reactor: 1 wt % Agitator RPM: 1000 Liquid Flow: 60.4ml/min Pressure: Variable (see data) Oxygen Flow Rate: Variable (seedata) Temperature: 100° C. Reaction Mass: 725 g Impeller Type: radial(1.25″)

TABLE 41 Oxidation Results for Example 33 Time Pres. O₂ Flow NPMIDAGlyphosate HCHO HCOOH min (psig) (sccm) (wt %) (wt %) (ppm) (ppm) 0 100225 1.52 5.72 951 2822 60 100 225 0.48 7.55 1127 4404 121 100 225 0.477.62 1219 4419 181 100 225 0.46 7.57 1272 4442 241 100 225 0.45 7.631301 4434 302 100 225 0.45 7.74 1351 4590 362 100 315 0.19 7.68 14674230 422 100 315 0.10 7.65 1518 3739 482 100 315 0.10 7.77 1633 3756 543100 315 0.10 7.77 1684 3714 603 100 315 0.11 7.75 1671 3741 663 99 3150.10 7.78 1724 3721 724 100 292 0.13 7.75 1706 3840 784 100 292 0.137.84 1758 3905 844 100 292 0.13 7.76 1748 3908 904 100 292 0.11 7.811608 3884 965 100 292 0.11 7.75 1659 3901 1025 100 292 0.11 7.91 17033973 1085 100 292 0.11 8.05 1819 4108 1145 100 270 0.15 7.82 1681 41201206 100 270 0.14 7.74 1687 4438 1266 100 270 0.16 7.52 1850 4063 1326100 270 0.13 7.42 1754 3962 1748 100 247 0.64 6.77 1566 4025 1809 100247 0.65 7.70 1572 4266 1869 100 247 0.65 7.52 1545 4313 1929 100 2470.67 7.54 1612 4473 1989 100 247 0.63 7.50 1620 4436 2050 100 247 0.597.59 1640 4500 2110 100 225 0.59 7.18 1562 4321 2170 100 225 0.68 7.501639 4517 2261 100 225 0.66 7.60 1659 4551 2351 100 225 0.59 7.41 16774519 2411 100 360 0.16 7.86 1907 4157 2472 100 360 0.16 7.86 2039 40402532 100 360 0.10 7.89 2028 3664 2592 100 360 0.09 7.98 2045 3635 2653100 360 0.10 7.82 2150 3628 2713 70 360 0.11 8.47 2036 4008 2773 71 3600.08 7.55 1338 2920

Example 34 Continuous Oxidation of NPMIDA to Glyphosate in the Presenceof a Pt/Fe/C Catalyst

This example demonstrates the continuous oxidation of NPMIDA toglyphosate in the presence of a Pt/Fe/C heterogeneous particulatecatalyst in a fixed bed reactor. The experiment was designed to mimic asmall initial section of a fixed bed reactor with gas and liquid feedsentering co-currently.

The experiment was performed in a continuous reactor system comprising avertical stainless steel tubular reactor (2.2 cm inside diameter; 61.5cm length; 215 ml volume). The gas and liquid feeds entered the tubularreactor at the top and flowed down through the reactor. The reactor wasfilled with a mixture of catalyst (50 g) and glass Raschig rings (6 mm).The catalyst comprised 2% by weight platinum and 0.2% by weight iron on⅛-inch carbon granule supports. The reactor was heated to about 90° C.with a heated water feed and brought to a pressure of about 100 psigwith nitrogen. After the reactor reached a temperature of 90° C., waterand nitrogen flow were stopped and the liquid feed and oxygen feed wereinitiated.

The liquid feed was fed to the top of the reactor at 90° C. andcomprised an aqueous slurry feed material containing NPMIDA (3.00% byweight) and formic acid (0.54% by weight). Oxygen was fed to the top ofthe reactor with the reactor pressure maintained at 100 psig. The liquidand oxygen feed rates were varied in a series of four experiments asindicated in Table 42 below. In each experiment, the system was allowedto equilibrate for at least one half hour before samples were collectedat the column exit and analyzed for formic acid and glyphosate.

TABLE 42 Effluent Analysis Under Different Operating Conditions forExample 34 Liquid Flow Oxygen Flow % % Formic % (ml/min) (sccm)Glyphosate Acid Formaldehyde 100 100 0.11 0.47 ~0.011 100 200 0.13 0.46~0.014 50 100 0.24 0.38 ~0.018 25 100 0.41 0.33 ~0.026

Example 35 Continuous Oxidation of NPMIDA to Glyphosate in Two StirredTank Reactors in Series with Catalyst Recycle and Crystallizer Recycle

This example demonstrates the continuous oxidation of NPMIDA toglyphosate in the presence of a heterogeneous particulate catalystslurry in a continuous reactor system comprising two stirred tankreactors staged in series. The reactor system was similar to that shownin FIG. 30. The two stirred tank reactors (R1 and R2) were as describedin Example 24 except that the impeller configuration of R2 was notoperated in a DISPERSIMAX mode. The catalyst was continuously filteredfrom the reaction mixture effluent withdrawn from R2 using a back-pulsefilter system comprising parallel filter bodies and the separatedcatalyst was recycled to R1. Crystalline N-(phosphonomethyl)glycineproduct was recovered from the filtrate in a crystallizer (30 L) and themother liquor from the crystallizer was recycled to R1.

The heterogeneous particulate catalyst was prepared by a method similarto that described in Example 17 above and comprised platinum (5% byweight) and iron (0.5% by weight) on a particulate carbon support. Inthis example, the particulate heterogeneous catalyst was transferredfrom R1 to R2 with the effluent from R1 including some entrained gas.The catalyst exited R2 with the reactor effluent, was separated in theback-pulse filter, and recycled back to R1. The back-pulse catalystfilter also acted as a liquid gas separator for the effluent from R2.The filtrate from the back-pulse catalyst filter was sent to acrystallizer for recovery of crystalline N-(phosphonomethyl)glycineproduct. The resulting mother liquor from the crystallizer was used toback-pulse the catalyst filter bodies and recycled to R1 with theseparated catalyst.

The operating conditions for R1 and R2 are summarized in Table 43. R1and R2 were charged initially as shown in Table 43 and oxygen wasintroduced concurrent with the NPMIDA feed. The NPMIDA feed comprised anaqueous slurry feed material containing from about 12.5% to about 15%NPMIDA and mother liquor recycle from the catalyst filters to give aneffective combined feed to R1. The effective combined feed to R1 wasintroduced initially at 4.3 wt % and later increased to 5.2 wt %.Bismuth oxide was added throughout the run to increase the formic aciddestruction rate. Bismuth oxide was added in a batchwise fashion to R1(˜5 mg per addition) and also in a continuous fashion by addition to theNPMIDA slurry feed (4-25 mg per 20 kg of NPMIDA slurry). The frequencyand amount of bismuth oxide added to the slurry feed is listed in Table44. The aqueous slurry feed to R1 (including the component from thecrystallizer mother liquor recycled with the catalyst), R1 reactoreffluent and R2 reactor effluent were analyzed by HPLC. The HPLCanalytical results are presented in Table 45.

TABLE 43 Operating Conditions for the Continuous Oxidation ReactorSystem of Example 35 Initial Reactor Charge R1 R2 Catalyst 0.8 wt % 0.8wt % NPMIDA 0.8 wt % 0.3 wt % Glyphosate 5.0 wt % 5.0 wt % formaldehyde500 ppm 500 ppm formic acid 2000 ppm 2000 ppm water 2700 ml 1500 mlOperating Conditions Catalyst Concentration¹ 0.8-1.4 wt % 0.8-1.4 wt %Agitator RPM 1000 600 Total Liquid Flow into R1 147.4 g/min 147.4 g/minPressure 100 psig 100 psig Oxygen 900-1700 sccm 120-700 sccm Temperature95°-100° C. 95°-100° C. Reaction Mass 2950 g 1725 g Impeller Type radial(2″)² radial (2″)² NPMIDA Slurry Flow Rate 50 g/min NA RML Flow Rate³97.4 ml/min NA ¹Initial catalyst charge was 0.8 wt %. During the run thecatalyst loading was increased to 1.0 wt % at 69 hours, 1.2 wt % at 119hours and 1.4 wt % at 143 hours. ²A downward pumping impeller wasinstalled on the agitator shaft about half way up the liquid column.³Crystallizer mother liquor (RML) was used to back pulse the catalystfilters and return filtered catalyst back to R1 with RML.

TABLE 44 Frequency and Amount of Bismuth Oxide Addition to ReactorSystem Elapsed Time (hrs) Bismuth Oxide Addition 19.4 12.5 mg to R1 with50 ml H₂O 19.6 4.2 mg to 20 kg of NPMIDA slurry feed 46.2 12.7 mg in R1with 50 ml H₂O 102.3 12.8 mg to R1 with 50 ml H₂O 4.7 mg to 20 kg ofNPMIDA slurry feed 139.3 12.4 mg to R1 with 50 ml H₂O 4.1 mg to 20 kg ofNPMIDA slurry feed 159.3 13.3 mg to R1 with 50 ml H₂O 4.0 mg to 20 kg ofNPMIDA slurry feed 165.7 16.4 mg to R1 with 50 ml H₂O 16.9 mg to 20 kgof NPMIDA slurry feed 172.0 25 mg to 20 kg of NPMIDA slurry feed 215.912.3 mg to 20 kg of NPMIDA slurry feed 370.7 16.7 mg to 20 kg of NPMIDAslurry feed

TABLE 45 Analytical Results for Example 35 Elapsed Effective Combined R1Feed R1 Exit R2 Exit time NPMIDA Gly HCHO HCOOH NPMIDA Gly HCHO HCOOHNPMIDA Gly HCHO HCOOH (hrs) (wt %) (wt %) (ppm) (ppm) (wt %) (wt %)(ppm) (ppm) (wt %) (wt %) (ppm) (ppm) 9.0 4.4 1.07 417 1775 0.62 3.86841 3058 0.54 3.96 449 3147 12.1 4.31 1.08 408 1857 0.36 3.88 927 34460.24 3.98 417 3448 16.9 4.30 1.12 497 2179 0.32 4.10 984 3507 0.18 4.13418 3416 18.0 4.30 1.03 492 2179 0.34 4.04 951 3664 0.19 3.80 401 341420.5 4.33 1.03 480 1853 0.57 4.35 1010 2941 0.29 3.81 354 2213 22.0 4.341.13 494 2082 0.52 4.20 943 2790 0.32 4.18 406 3058 27.0 4.31 2.27 2691368 0.45 4.62 1035 3147 0.23 4.53 445 2911 30.0 4.32 2.54 310 1491 0.485.31 1196 3331 0.26 5.54 596 3364 30.7 4.31 2.57 311 1646 0.44 5.34 11783369 0.21 5.61 567 3347 34.5 4.27 2.43 406 1882 0.12 5.51 1282 3156 0.095.41 824 3364 36.0 4.29 2.42 432 1887 0.18 5.37 1332 3299 0.14 5.40 9203381 39.0 4.30 2.44 459 1969 0.19 5.67 1049 3704 0.19 5.45 1017 368446.0 4.36 2.55 483 2176 0.52 5.65 1660 4224 0.41 5.88 1106 4447 47.34.42 2.46 429 1677 0.74 5.88 1567 3528 0.62 5.55 909 2608 48.3 4.39 2.44405 1864 0.64 5.44 1382 2809 0.54 5.47 819 3297 50.7 4.44 2.76 552 22250.45 5.50 1505 3417 0.26 5.61 805 3245 56.5 4.35 2.59 518 1827 0.09 5.351566 2807 0.00 5.64 724 2113 59.0 4.37 2.51 591 2165 0.15 5.21 1700 33680.06 5.32 993 3359 62.1 4.32 2.51 649 2284 0.22 5.18 1760 3495 0.10 5.361087 3520 65.0 4.29 2.54 704 2213 0.17 5.22 1890 3634 0.06 5.42 11783718 69.5 5.18 2.53 577 1805 0.21 5.42 1402 2211 0.19 5.38 710 2215 71.55.20 2.65 600 1898 0.35 5.74 1574 2468 0.22 5.88 881 2735 75.0 5.27 2.74612 2134 0.56 5.08 1577 3240 0.40 6.06 958 3704 77.5 5.25 2.74 612 18280.44 5.84 1666 3256 0.27 6.06 952 2503 80.8 5.25 2.63 616 1802 0.46 5.751717 3406 0.26 5.90 961 2436 84.0 5.27 2.85 663 2114 0.50 5.84 1680 33280.31 6.47 1089 3437 87.0 5.27 2.63 656 2004 0.46 5.92 1735 3288 0.295.56 979 2897 91.0 5.26 2.47 687 1933 0.45 6.03 1924 3173 0.23 5.37 10122906 93.0 5.26 2.74 752 2122 0.45 6.10 1967 3706 0.23 6.37 1251 360597.7 5.24 2.69 898 2684 0.45 5.94 2248 4651 0.19 6.20 1569 4428 105.05.23 2.61 968 2210 0.42 5.79 2195 3791 0.21 6.11 1478 3276 107.0 5.242.64 1001 2388 0.48 5.77 2265 3935 0.24 5.95 1558 3403 109.0 5.24 2.631031 2242 0.52 5.80 2255 3697 0.26 6.16 1629 3480 111.0 5.24 2.59 10312242 0.52 5.80 2255 3697 0.26 6.03 1629 3480 116.0 5.19 2.68 1213 26200.46 6.02 2634 4379 0.14 6.15 1976 4163 129.5 5.27 2.66 1328 3239 0.705.91 2952 6163 0.42 6.05 2400 6449 134.0 5.27 2.58 1332 2880 0.76 5.723014 5731 0.43 5.78 2415 5125 135.5 5.41 2.65 1310 2872 1.14 5.62 28484719 0.96 6.02 2336 5095 140.3 5.23 2.84 1268 3477 0.52 6.34 3494 56130.15 6.53 2604 5308 142.0 5.21 2.78 1248 3364 0.39 6.07 3429 5113 0.096.30 2529 4891 144.6 5.20 2.84 849 3154 0.25 6.29 1751 4150 0.03 6.521058 4117 152.5 5.23 2.62 1026 3065 0.34 5.67 2330 3998 0.16 5.70 17113788 153.5 5.19 2.63 1046 2580 0.37 5.67 2317 4048 0.14 5.75 1724 3726156.5 5.19 2.72 1135 2962 0.46 5.84 2729 5120 0.16 6.07 2011 5137 158.35.19 2.61 1168 2920 0.48 5.85 2747 5363 0.14 5.66 2132 4980 160.5 5.242.71 1150 2696 0.63 5.86 2750 5035 0.32 6.06 2066 4156 161.5 5.19 2.781135 2729 0.51 6.09 2670 4389 0.15 6.32 2010 4277 163.5 5.18 2.89 11923068 0.54 6.13 2844 5696 0.10 6.73 2219 5528 167.1 5.22 2.73 1261 32500.67 5.50 2659 5220 0.28 6.14 2476 6197 170.0 5.14 2.77 1431 3625 0.455.93 3133 5555 0.03 6.45 2474 5485 173.0 5.12 2.64 1484 3346 0.28 5.273099 4459 0.01 5.73 2564 4327 176.3 5.15 2.68 1528 3489 0.56 5.50 31755094 0.12 5.95 2649 5136 179.8 5.15 2.74 1280 3225 0.55 5.78 3222 47160.11 6.16 2587 4164 183.5 5.17 2.73 1306 3247 0.74 5.93 3308 4556 0.196.13 2684 4245 188.0 5.15 2.64 1244 3043 0.54 5.61 3143 3924 0.14 5.782455 3491 189.5 5.14 2.60 1624 2184 0.37 5.97 3165 3278 0.09 6.12 24813057 190.5 5.14 2.62 1601 1858 0.38 5.99 3109 3059 0.09 6.03 2417 2358196.0 5.14 2.60 1624 2184 0.37 5.97 3165 3278 0.09 6.12 2481 3057 199.05.14 2.62 1601 1858 0.38 5.99 3109 3059 0.09 6.03 2417 2358 201.5 5.142.70 1559 1600 0.28 5.74 2889 2528 0.08 6.24 2305 2110 203.5 5.13 2.651525 1522 0.29 6.01 3001 2260 0.07 6.11 2241 1898 205.5 5.14 2.63 15171439 0.30 5.95 2871 2104 0.10 6.05 2210 1590 207.5 5.13 2.90 1448 11300.23 5.95 2819 1786 0.06 6.15 2065 1319 209.5 5.14 2.77 1492 1010 0.286.10 2929 1839 0.10 6.34 2224 1275 211.5 5.14 2.69 1453 932 0.27 5.642774 1736 0.10 6.05 2082 986 219.5 5.14 2.64 1287 585 0.28 5.81 25541234 0.07 6.02 1858 819 221.5 5.14 2.91 1273 538 0.37 5.86 2605 12270.10 6.03 1860 748 223.5 5.15 2.64 1272 530 0.35 5.95 2558 1146 0.095.97 1846 751 225.5 5.16 2.71 1262 489 0.42 5.98 2573 1144 0.12 6.121869 688 227.5 5.15 2.79 1254 473 0.39 5.78 2500 1122 0.10 6.40 1841 629229.5 5.12 2.58 1222 165 0.36 5.95 2505 1028 0.08 5.70 1806 608 231.55.15 2.82 1222 429 0.39 6.09 2550 948 0.11 6.36 1844 595 233.5 5.16 2.811214 425 0.40 6.05 2514 902 0.12 6.32 1816 581 242.5 5.16 2.64 1131 3260.51 5.91 2372 867 0.09 5.85 1597 495 244.0 5.16 2.50 1105 313 0.56 5.982416 791 0.11 5.36 1499 450 245.5 5.20 2.82 1186 314 0.71 6.07 2441 7910.25 6.43 1796 469 248.0 5.14 2.81 1150 303 0.48 5.95 2458 746 0.03 6.411660 432 249.3 5.16 2.81 1165 304 0.45 5.74 2310 699 0.10 6.41 1717 432250.3 5.17 2.79 1162 303 0.52 6.10 2409 787 0.13 6.34 1707 429 251.55.18 2.75 1150 287 0.58 5.72 2250 701 0.16 6.17 1663 370 256.0 5.16 2.881101 254 0.60 5.66 2166 625 0.03 6.43 1619 391 258.0 5.17 2.87 1108 3100.81 5.91 2239 651 0.11 6.50 1693 429 260.0 5.16 2.83 1093 300 0.73 5.782234 612 0.06 6.35 1638 394 262.0 5.15 2.79 1091 249 0.71 5.77 2270 6400.07 6.41 1674 395 267.0 5.30 2.70 1062 213 1.33 5.29 2049 531 0.55 6.001645 315 270.0 5.25 2.84 1080 222 1.12 5.88 2226 587 0.35 6.51 1710 349271.0 5.20 2.87 1069 224 0.95 5.91 2243 685 0.19 6.60 1669 358 273.25.21 2.87 1071 247 0.90 5.93 2207 586 0.20 6.61 1675 440 273.8 5.21 2.831076 231 0.90 5.82 2235 640 0.23 6.48 1696 383 275.5 5.21 2.90 1107 2370.93 5.68 2167 596 0.20 6.72 1809 406 277.0 5.18 2.87 1100 270 0.91 6.282240 590 0.12 6.61 1785 527 278.2 5.19 2.89 1163 240 0.94 5.95 2339 6320.11 6.76 1852 422 281.5 5.20 2.79 1186 281 0.92 6.08 2505 729 0.12 6.391936 571 284.5 5.19 2.71 1162 244 0.87 6.04 2630 708 0.07 6.09 1847 437287.5 5.19 2.69 1179 269 0.84 6.02 2693 745 0.07 6.03 1908 529 292.55.13 2.64 1177 241 0.20 5.87 2679 653 0.07 5.96 1718 378 301.0 5.27 2.65979 166 0.92 5.30 1839 406 0.45 5.86 1259 200 335.5 5.50 1.14 191 821.85 3.64 1200 505 1.48 4.21 703 302 339.8 5.10 1.34 226 74 0.15 4.841654 518 0.01 4.92 833 273 344.0 5.11 1.41 235 38 0.48 5.25 1824 4800.04 5.20 868 142 347.5 5.12 1.44 285 68 0.33 4.96 1867 528 0.09 5.301050 252 351.5 5.12 1.42 297 69 0.28 5.18 2003 474 0.07 5.24 1096 255353.8 5.11 1.44 314 75 0.30 5.11 2115 753 0.05 5.29 1157 276 357.5 5.162.94 807 216 0.44 6.34 2417 757 0.16 6.74 1558 364 359.5 5.13 2.83 789192 0.29 6.27 2553 943 0.04 6.34 1490 278 368.0 5.15 2.70 881 201 0.345.96 2792 694 0.13 5.85 1831 309 370.5 5.14 2.91 894 227 0.34 6.65 2897950 0.09 6.64 1877 406 373.5 5.17 2.95 986 224 0.51 6.49 3084 542 0.196.77 2216 394 375.0 5.17 2.78 1425 270 0.43 6.46 3131 976 0.12 6.44 2291563 378.3 5.14 2.76 1437 238 0.32 6.05 3273 940 0.03 6.34 2334 444 380.55.14 2.79 1425 226 0.21 6.13 3229 1021 0.00 6.48 2290 402 383.5 5.342.87 1853 272 0.95 6.57 3829 1031 0.59 6.55 3052 397 387.5 5.20 2.861868 275 0.25 6.09 3652 842 0.07 6.55 2787 476 389.5 5.22 2.82 1953 3080.53 5.89 3833 1370 0.22 6.43 3115 476 391.0 5.28 2.91 2002 334 0.726.05 4118 1358 0.42 6.74 3295 574 392.0 5.26 2.91 2061 286 0.59 6.414047 1336 0.32 6.73 3273 485 395.0 5.19 2.83 2101 324 0.31 6.19 40221036 0.10 6.62 3221 582 398.0 5.16 2.76 2153 344 0.34 6.17 4219 15190.09 6.47 3353 649 399.5 5.16 2.78 2184 396 0.33 6.13 4262 1286 0.076.53 3466 840 401.5 5.15 2.78 2265 494 0.38 6.19 4323 1570 0.09 6.373609 1005 405.0 5.21 2.88 2383 563 0.65 6.21 4337 1470 0.27 6.63 3731937 407.0 5.20 2.87 2409 635 0.61 6.30 4492 1922 0.25 6.58 3830 1202409.6 5.20 2.70 2493 708 0.53 6.40 4498 1926 0.17 5.84 3854 1360 412.05.14 2.66 2438 723 0.27 5.66 4230 1874 0.02 6.04 3601 1387 413.5 5.152.67 2497 747 0.36 5.71 4355 2236 0.07 6.09 3819 1477 416.5 5.13 2.682553 985 0.39 5.78 4371 2351 0.06 6.17 3911 1804

Example 36 Continuous Oxidation of NPMIDA to Glyphosate in Two StirredTank Reactors in Series with Catalyst Recycle and Crystallizer Recycle

This example demonstrates the continuous oxidation of NPMIDA toN-(phosphonomethyl)glycine in a continuous reactor system comprising twostirred-tank reactors staged in series utilizing a heterogeneousparticulate catalyst slurry. The reactor system was similar to thatshown in FIG. 31. The two stirred tank reactors (R1 and R2) were asdescribed in Example 35 above. The catalyst was continuously filteredfrom the reaction mixture effluent withdrawn from R2 using a back-pulsefilter system comprising parallel filter bodies and the separatedcatalyst was recycled to R1. Crystalline N-(phosphonomethyl)glycineproduct was recovered from the filtrate in a crystallizer (30 L) and themother liquor from the crystallizer was recycled to R1. Additionally, aportion of the crystallizer mother liquor was added to the effluent fromR2 as an effluent dilution stream (75-100 ml/min) to reduce theglyphosate concentration in the R2 effluent introduced to the back-pulsefilter to reduce potential crystallization problems. Also, the reactionsystem further comprised a catalyst rest tank (500 ml Hastelloy Cautoclave with an upward pumping impeller), which collected theseparated catalyst prior to its re-introduction to R1. The catalyst resttank was operated without level control and catalyst slurry was allowedto exit at the top of the vessel.

The heterogeneous particulate catalyst was prepared by a method similarto that described in Example 17 above and comprised platinum (5% byweight) and iron (0.5% by weight) on a particulate carbon support. Inthis example, the particulate heterogeneous catalyst was transferredfrom R1 to R2 with the effluent from R1 including some entrained gas.The catalyst exited R2 with the reactor effluent, was separated in theback-pulse filter, sent to the catalyst rest tank and recycled back toR1. The back-pulse catalyst filter also acted as a liquid gas separatorfor the effluent from R2. The filtrate from the back-pulse catalystfilter was sent to a crystallizer for recovery of crystallineN-(phosphonomethyl)glycine product. The resulting mother liquor from thecrystallizer was used to back-pulse the catalyst filter bodies, as adiluent for the effluent from R2 passing to the catalyst filter bodies,and recycled to R1 with the separated catalyst.

The operating conditions for R1 and R2 are summarized in Table 46. R1and R2 were charged initially as shown in Table 46 and oxygen wasintroduced concurrent with the NPMIDA feed. The NPMIDA feed comprised anaqueous slurry feed material containing from about 20% to about 20.5%NPMIDA and mother liquor recycle from the catalyst filters to give aneffective combined feed to R1. The effective combined feed to R1 wasintroduced initially at 7 wt % and later increased to 7.7 wt %. Bismuthoxide was added throughout the run to increase the formic aciddestruction rate. Bismuth oxide was added in a continuous fashion byaddition to the NPMIDA slurry feed (3-12 mg per 20 kg of NPMIDA slurry).The frequency and amount of bismuth oxide added to the slurry feed islisted in Table 47. The aqueous slurry feed to R1 (including thecomponent from the crystallizer mother liquor recycled with thecatalyst), R1 reactor effluent and R2 reactor effluent were analyzed byHPLC. The HPLC analytical results are presented in Table 48.

TABLE 46 Operating Conditions for Example 36. Initial Reactor Charge R1R2 Catalyst 1.6 wt % 1.6 wt % water 2400 ml 2400 ml Operating ConditionsCatalyst Concentration¹ 1.6-2.2 wt % 1.6-2.2 wt % Agitator RPM 1000 1000Total Liquid Flow into R1 147.4 g/min 147.4 g/min Pressure 100 psig 100psig Oxygen 1200-2750 sccm 350-1200 sccm Temperature 105°-110° C.105°-110° C. Reaction Mass 2950 g 2950 g Impeller Type radial (2″)²radial (2″)² NPMIDA Slurry Flow Rate 50 g/min NA RML Flow Rate³ 97.4ml/min NA ¹Initial catalyst charge was 1.6 wt %. During the run thecatalyst loading was increased to 1.7 wt % at 344 hours, 1.8 wt % at 354hours, 1.9 wt % at 356 hours, 2.0 wt % at 359 hours, 2.1 wt % at 363hours and 2.2 wt % at 366 hours. ²A downward pumping impeller wasinstalled on the agitator shaft about half way up the liquid column.³Crystallizer mother liquor (RML) was used to back pulse the catalystfilters and return filtered catalyst back to R1 with RML.

TABLE 47 Frequency and Amount of Bismuth Oxide Addition to NPMIDA SlurryFeed. Elapsed Time Bismuth Oxide Addition (hrs) (to 20 kg of NPMIDASlurry Feed) 9.8 3 mg 20.5 6 mg 34.3 9 mg 44.0 18 mg  89.4 12 mg  98.0 3mg 204.9 6 mg 225.1 3 mg 274.3 12 mg 

TABLE 48 Analytical Results for Example 36. Elapsed Effective CombinedR1 Feed R1 Exit R2 Exit time NPMIDA Gly HCHO HCOOH NPMIDA Gly HCHO HCOOHNPMIDA Gly HCHO HCOOH (hrs) (wt %) (wt %) (ppm) (ppm) (wt %) (wt %)(ppm) (ppm) (wt %) (wt %) (ppm) (ppm) 2.2 8.03 1.84 185 1338 3.50 4.47682 3378 3.03 4.54 456 3303 5.5 6.85 1.71 53 515 1.00 6.14 712 3816 0.206.88 215 2073 8.5 6.86 2.55 71 742 1.10 6.33 743 3742 0.22 6.84 230 256411.5 6.85 2.85 97 1044 1.01 7.05 842 4090 0.17 7.56 247 2609 14.5 6.842.28 95 1079 0.64 6.89 817 3927 0.12 7.48 216 2398 17.5 7.01 2.67 1461441 1.44 6.38 849 4142 0.74 7.06 399 3494 20.5 7.12 2.72 188 1746 1.606.22 837 4323 0.70 6.86 388 3458 23.5 7.13 2.87 213 2115 1.30 6.82 8924732 0.49 7.43 385 3674 26.5 7.01 3.00 201 1854 1.06 7.39 904 4871 0.137.67 298 2414 28.0 6.99 2.93 193 1827 1.04 6.79 951 4147 0.06 7.38 2672303 30.0 7.00 2.96 204 1941 1.01 6.88 973 4621 0.10 7.50 311 2779 31.06.92 2.80 215 1888 1.31 6.81 1022 6357 0.18 7.57 385 3050 33.0 6.89 2.76217 1721 1.09 6.45 975 4210 0.13 7.36 345 2629 35.0 6.87 2.75 226 18880.99 6.78 1198 4748 0.06 7.50 392 2761 36.5 6.87 2.66 270 1885 0.94 6.671422 4872 0.05 7.10 576 2747 37.5 6.87 2.78 299 2138 0.88 6.92 1617 54030.05 7.62 694 3804 40.0 7.73 2.72 423 2844 1.35 6.84 1867 6272 0.21 6.99727 5318 43.0 7.75 2.87 419 3330 1.45 6.78 1871 6615 0.20 7.57 696 567244.8 7.75 2.86 726 3316 1.28 6.69 2207 6247 0.19 7.50 1262 5180 46.37.77 2.79 761 3106 1.72 6.99 2216 6140 0.29 7.20 1405 4303 48.0 7.792.89 768 3357 1.44 7.00 2309 6196 0.35 7.64 1437 5353 51.0 7.84 3.691033 4607 1.22 7.16 2348 5807 0.16 7.92 1379 4709 53.0 7.83 3.64 10024423 1.29 6.66 2323 5481 0.09 7.70 1249 3940 54.8 7.94 3.46 1066 45131.94 8.07 2234 5273 0.57 7.24 1497 3870 56.0 7.90 3.00 933 3216 2.046.92 2304 5384 0.57 7.84 1557 4666 57.0 7.85 2.89 915 2927 1.60 7.192307 5030 0.36 7.34 1473 3327 58.5 7.86 3.14 941 3065 1.81 7.14 23575248 0.34 8.04 1470 4184 60.5 7.85 3.40 940 3010 1.35 7.30 2399 48710.30 9.24 1468 3927 62.5 7.81 3.05 923 2729 1.34 7.34 2448 4730 0.168.00 1349 3329 64.0 7.82 3.02 934 2564 1.69 8.50 2420 4719 0.20 7.871401 2564 65.5 7.87 3.22 986 2829 1.63 7.69 2364 4381 0.42 8.80 16423791 67.0 7.78 2.75 866 2291 0.51 6.94 2377 3248 0.02 6.64 1085 129368.0 7.79 2.96 910 2616 0.99 7.33 2498 3884 0.09 7.58 1289 2801 69.07.81 3.00 920 2476 1.16 6.93 2543 3938 0.15 7.78 1334 2151 71.0 7.743.10 890 1969 1.18 7.18 2528 3862 0.12 8.12 1415 2796 73.3 7.75 3.14 9061988 1.46 7.73 2547 4093 0.18 8.31 1493 2881 75.5 7.75 3.14 953 19601.53 7.15 2398 3595 0.17 8.29 1534 2773 77.8 7.73 3.12 933 1994 1.217.28 2424 3353 0.08 8.23 1441 2932 79.5 7.71 3.17 940 1725 1.29 7.162448 3237 0.07 7.96 1436 2264 81.0 7.72 3.01 937 1558 1.41 7.95 24382738 0.10 7.23 1421 1490 82.5 7.73 3.17 997 1709 1.46 7.07 2445 30900.13 8.12 1590 2294 84.0 7.73 3.07 1003 1546 1.49 7.89 2459 2579 0.157.67 1616 1536 86.8 7.73 3.11 1010 1509 1.57 8.16 2497 2442 0.11 7.841648 1364 88.5 7.72 3.27 1092 1500 1.46 7.44 2608 2896 0.09 8.57 17392048 90.0 7.75 3.48 1109 1518 2.04 7.73 2696 2969 0.21 9.50 1818 212991.5 7.75 3.08 1159 1529 1.91 7.44 2671 2935 0.20 7.86 1907 2219 93.07.74 3.23 1152 1340 1.74 8.07 2971 2964 0.14 8.54 1874 1342 94.5 7.733.71 1192 1425 1.56 7.77 3039 2883 0.10 9.01 1893 1899 98.5 7.66 1.67264 233 1.28 6.26 2256 1985 0.05 7.73 1227 1082 100.0 7.65 1.54 160 1661.02 6.42 1807 1496 0.02 7.16 742 769 101.5 7.65 1.54 146 153 0.97 6.581718 1927 0.01 7.13 678 709 104.5 7.65 1.45 122 151 0.71 6.01 1591 17390.02 6.73 566 699 107.5 7.68 1.42 118 175 1.10 5.71 1439 1698 0.14 6.59549 811 111.5 7.75 1.54 130 196 1.39 6.20 1429 1854 0.46 7.15 603 909113.0 7.75 2.70 190 164 1.23 5.67 1337 1717 0.08 6.71 520 479 114.2 7.843.10 237 310 1.98 7.39 1605 2358 0.48 8.55 740 1156 115.5 7.75 3.23 305349 1.15 7.32 1662 2044 0.07 8.32 694 833 118.5 7.77 3.25 288 243 1.507.49 1701 1973 0.21 8.99 750 647 121.5 7.79 3.48 555 756 1.56 7.82 21112604 0.28 8.81 1093 1395 124.5 7.73 3.09 507 677 1.46 8.01 2239 27570.16 9.04 1122 1404 127.5 7.82 3.61 719 1064 1.57 7.64 2283 2755 0.498.99 1249 1530 133.5 7.71 3.22 835 1021 1.53 7.75 2507 2738 0.11 8.551400 1583 136.5 7.70 3.40 906 1090 1.29 7.78 2634 2961 0.09 8.72 14941626 139.5 7.77 3.55 1124 1181 1.48 6.26 2718 1929 0.36 8.74 2267 2129142.5 7.88 3.50 1045 1064 1.68 7.72 2733 2985 0.54 8.75 1690 1569 145.57.77 3.60 1055 1194 1.53 8.07 2836 2997 0.07 9.05 1655 1557 149.0 7.723.52 1100 1359 1.32 7.62 2949 2913 0.06 8.81 1773 1666 152.0 7.71 3.341138 1302 1.25 7.62 3032 2408 0.11 8.79 1811 1602 155.0 7.72 3.62 11651354 1.79 8.32 2989 3128 0.19 9.16 1864 1584 158.5 7.71 3.98 1195 11641.35 8.02 3162 2902 0.10 8.92 1933 1641 161.5 7.70 3.47 1276 1193 1.347.75 3174 3005 0.12 8.89 2071 1691 164.5 7.73 3.13 1324 1059 1.40 7.823264 2969 0.13 7.42 2135 1263 167.5 7.71 3.38 1311 918 1.47 7.99 33783102 0.12 8.61 2089 1528 172.0 7.74 2.93 1427 1173 1.44 7.50 3428 27310.14 8.31 2223 1511 175.0 7.72 3.07 1480 1050 1.48 7.51 3555 2653 0.108.57 2314 1479 178.0 7.69 2.91 1474 1159 1.39 6.89 3347 2254 0.09 8.482319 1506 181.0 7.69 2.91 1505 1044 1.42 7.79 3573 2669 0.10 8.43 23221452 184.0 7.70 3.01 1525 1058 1.39 7.41 3572 2670 0.11 8.81 2268 1499186.5 7.78 2.81 1802 2335 2.69 6.32 3714 6849 0.51 7.87 3550 7428 188.38.02 2.47 1418 3234 4.05 6.45 3407 8082 1.24 7.99 3294 8694 203.5 7.762.59 1303 2043 1.76 6.65 3365 4733 0.15 7.53 2552 3664 205.8 7.79 2.741397 2401 2.04 7.25 3564 5706 0.29 8.20 2988 5327 209.5 7.66 1.53 544523 0.91 6.13 3720 4459 0.04 7.09 2525 2427 212.0 7.72 2.60 1454 17881.51 6.92 3834 4230 0.11 7.55 2802 2741 213.5 7.74 2.61 1439 1954 2.247.04 3462 4683 0.19 7.59 2732 3513 215.0 7.75 2.72 1819 2043 1.85 7.014071 4994 0.22 7.03 3126 3839 216.5 7.76 2.86 1769 2046 1.93 7.02 35774998 0.25 7.70 2892 3852 218.0 7.74 2.75 1959 2623 1.50 6.88 4035 51010.13 8.13 3122 3909 221.0 7.73 2.77 2080 2861 1.68 6.91 4039 5394 0.157.94 3284 4107 225.0 7.72 2.96 1808 2461 1.59 6.99 3427 5187 0.14 7.432706 3819 228.5 7.65 1.44 535 407 0.76 6.99 3674 2202 0.03 6.71 24841888 232.0 7.68 2.54 1208 524 1.12 7.01 2708 2686 0.13 7.24 1619 1724236.5 7.70 2.98 2767 1236 1.20 7.53 2934 3346 0.11 8.46 1820 2154 238.57.71 2.95 2962 1256 1.51 7.51 3748 3631 0.17 8.31 2728 2246 242.0 7.732.93 1404 1146 1.48 7.14 3231 3767 0.17 7.55 2192 2064 245.5 7.74 3.001778 1657 1.44 7.19 3758 3684 0.13 8.22 2782 2247 248.0 7.72 3.00 18311739 1.47 7.24 3781 3763 0.14 8.39 2867 2397 250.0 7.72 2.96 1834 17331.57 8.02 3875 3928 0.15 8.20 2882 2372 252.3 7.69 3.08 1995 1421 1.477.53 4141 4229 0.13 8.50 3190 2405 255.0 7.70 3.17 2045 1699 1.60 7.644102 3635 0.15 8.74 3076 2601 258.0 7.72 2.92 1919 1898 1.57 7.39 39144148 0.14 8.08 2778 2428 261.0 7.72 2.98 1909 2016 1.54 7.24 3935 41090.13 8.09 2747 2603 262.5 7.73 2.97 1853 2044 1.76 7.30 3296 4404 0.158.06 2490 2731 264.0 7.73 2.92 1905 2125 1.83 7.41 3858 4367 0.16 8.222720 2828 270.0 7.73 2.93 1978 2269 1.86 7.15 3840 4415 0.15 8.30 30623498 273.0 7.71 3.01 1646 1713 1.81 6.97 3081 4619 0.13 7.97 2492 3091276.0 7.72 2.95 1637 1751 1.87 6.95 3062 4531 0.16 7.99 2471 3038 279.07.72 2.94 1646 2020 1.95 6.90 3049 4576 0.14 7.99 2514 2600 282.0 7.742.99 1973 1741 1.67 7.02 3093 3462 0.16 8.05 2450 2548 285.0 7.69 2.961807 1438 1.53 7.23 3444 3660 0.08 8.11 2547 2352 288.0 7.70 2.98 17811464 1.46 7.13 3498 3683 0.10 7.97 2654 2359 291.0 7.69 2.92 1809 13711.55 7.10 3520 3793 0.11 8.03 2756 2450 294.0 7.69 2.89 1823 1357 1.527.13 3666 3750 0.10 7.96 2790 2399 297.0 7.70 2.88 1836 1377 1.82 6.983632 3747 0.14 7.92 2849 2492 300.0 7.69 2.89 1828 1358 1.71 6.85 35253640 0.11 7.95 2815 2406 303.0 7.71 2.86 1889 1468 1.93 6.87 3617 38020.14 7.78 2868 2505 306.0 7.71 2.85 1875 1453 1.72 6.65 3538 3657 0.117.71 2803 2435 309.0 7.68 2.80 1874 1356 1.60 6.70 3573 3731 0.09 7.722904 2512 312.0 7.70 2.80 2117 2129 1.96 6.83 3902 3753 0.13 8.04 32232603 315.0 7.70 2.89 2106 1744 1.85 6.94 3946 3545 0.10 8.96 3459 2368318.0 7.72 2.85 2115 2025 2.21 6.89 3933 3772 0.16 8.23 3299 2731 321.07.72 2.84 2150 2177 2.24 6.73 3864 3918 0.16 8.20 3313 2989 324.0 7.732.82 2173 2207 2.21 6.69 3857 3817 0.15 8.27 3321 2920 327.0 7.73 2.942187 2244 2.27 6.69 3805 4088 0.17 8.19 3317 3221 330.0 7.74 2.95 23821790 2.46 6.76 4069 4402 0.23 8.14 3536 3297 333.0 7.78 3.05 2777 19422.52 6.67 4864 4586 0.21 8.48 4232 3341 337.0 7.75 3.07 2848 1987 2.197.20 5059 4908 0.15 8.35 4237 3434 339.5 7.75 2.97 2798 2025 2.35 6.664770 4719 0.20 8.06 4152 3552 342.5 7.75 3.07 2782 2105 2.39 6.98 49054929 0.18 8.36 4240 3636 344.0 7.75 3.07 2776 2564 2.37 6.90 4860 52420.17 8.43 4198 3899 346.0 7.71 2.97 2462 2316 1.56 7.34 4418 4337 0.088.16 3386 2732 348.0 7.72 3.01 2506 2515 1.98 7.00 4454 5019 0.12 8.343593 3655 350.3 7.71 2.97 2386 2328 1.90 7.19 4624 4795 0.11 8.17 36143321 352.3 7.71 3.00 2441 2387 1.95 6.96 4594 4764 0.13 8.27 3869 3596354.3 7.70 3.05 2531 1888 2.13 7.03 4941 4936 0.14 8.27 3839 3312 356.37.68 2.97 2289 1874 1.39 7.62 4250 4101 0.06 8.53 3191 2754 358.3 7.673.04 2021 1415 1.05 8.15 4080 3814 0.03 8.49 2685 2026 365.0 7.68 3.481358 998 1.13 8.47 3195 3102 0.02 7.95 1949 1597 367.0 7.65 3.91 12661063 1.13 8.33 2825 3149 0.02 8.94 1612 1647 368.0 7.65 3.87 1217 11011.24 8.31 2475 3257 0.02 9.08 1737 1898

Example 37 Continuous Oxidation of NPMIDA to Glyphosate in a Fixed BedReactor Using a Pt/Fe/C Catalyst

This example demonstrates the continuous oxidation of NPMIDA toglyphosate in the presence of a Pt/Fe/C heterogeneous particulatecatalyst in a fixed bed reactor. The experiment was designed to simulatea small section of a fixed bed reactor with gas and liquid feedsentering co-currently.

The experiment was conducted in a continuous reactor system comprising avertical stainless steel tubular reactor (1.56 cm ID, 60.5 cm in length,116 ml volume). The gas and liquid feed streams entered the tubularreactor at the top and flowed down through the reactor. The reactorcontained a Pt/Fe/C catalyst (42.3 g) comprising platinum (2% by weight)and iron (0.2% by weight) on 1.2 mm-diameter extruded carbon supportsranging from about 1 mm to about 9 mm in length. The reactor was heatedto 90° C. with a heated water feed and brought to a pressure of 150 psigwith nitrogen. After the reactor temperature reached 90° C., the waterand nitrogen flow were stopped and the liquid and oxygen feeds wereinitiated.

The liquid feed (50 ml/min) comprised an aqueous slurry feed materialcontaining NPMIDA (1.92% by weight), glyphosate (1.89% by weight),formic acid (0.26% by weight) and formaldehyde (0.15% by weight). Oxygen(200 sccm) was fed to the top of the reactor with the pressuremaintained at 150 psig. After ten days of continuous operation, analysisof the reactor product showed 0.15% formaldehyde, 0.26% formic acid and2.53% glyphosate.

Example 38 Continuous Oxidation of NPMIDA to Glyphosate in a Fixed BedReactor Using a Pt/Fe/C Catalyst

This example demonstrates the continuous oxidation of NPMIDA toglyphosate in the presence of a Pt/Fe/C heterogeneous particulatecatalyst in a fixed bed reactor with cocurrent upflow of gas and liquidreactants.

The experiment was conducted in a continuous oxidation reactor systemcomprising a vertical stainless steel tubular reactor as described inExample 37, except that reactants flowed up through the reaction zone.The reactor contained a Pt/Fe/C catalyst (42.3 g) comprising platinum(2% by weight) and iron (0.2% by weight) on 1.2 mm-diameter extrudedcarbon supports ranging from about 1 mm to about 9 mm in length. Thereactor was heated to 90° C. with a heated water feed and brought to apressure of 150 psig with nitrogen. After the reactor temperaturereached 90° C., the water and nitrogen flow were stopped and the liquidand oxygen feeds were initiated.

The liquid feed (50 ml/min) comprised an aqueous slurry feed materialcontaining NPMIDA (1.80% by weight), glyphosate (2.19% by weight),formic acid (0.26% by weight) and formaldehyde (0.14% by weight) and wasfed to the bottom of the reactor at 90° C. Oxygen (200 sccm) was fed tothe bottom of the reactor with the pressure maintained at 150 psig.After nineteen hours of continuous operation, analysis of the reactorproduct showed 0.13% formaldehyde, 0.16% formic acid and 2.42%glyphosate.

Example 39 Continuous Oxidation of NPMIDA Potassium Salt to GlyphosatePotassium Salt in a Fixed Bed Reactor Using a Pt/Fe/C Catalyst withCo-Current Upflow of Liquid and Gas Reactants

This example demonstrates the continuous oxidation of NPMIDA potassiumsalt to glyphosate potassium salt in the presence of a Pt/Fe/Cheterogeneous particulate catalyst in a fixed bed reactor withco-current upflow of liquid and gas reactants.

The experiment was conducted in a continuous oxidation reactor systemcomprising a vertical stainless steel tubular reactor as described inExample 37, except that reactants flowed up through the reaction zone.The reactor contained a Pt/Fe/C catalyst (42.3 g) comprising platinum(2% by weight) and iron (0.2% by weight) on 1.2 mm-diameter extrudedcarbon supports ranging from about 1 mm to about 9 mm in length. Thereactor was heated to 90° C. with a heated water feed and brought to apressure of 150 psig with nitrogen. After reaching 90° C., the water andnitrogen flow were stopped and the liquid and oxygen feeds wereinitiated.

The liquid feed (50 ml/min) comprised an aqueous slurry feed materialcontaining NPMIDA as the potassium salt (22.9% by weight), glyphosate(0.09% by weight), formic acid (0.20% by weight) and formaldehyde (0.14%by weight), and was fed to the bottom of the reactor at 90° C. Oxygen(500 sccm) was fed to the bottom of the reactor with the pressuremaintained at 150 psig. Analysis of the reactor product showed 0.35%formaldehyde, 0.20% formic acid and 1.56% glyphosate potassium salt.

Example 40 Comparison of Pt/Fe Catalyst Versus a Mixture of Pt/Fe andPt/Fe/Te Catalysts

This example compares the conversion of NPMIDA to glyphosate in acontinuous oxidation reactor system using a Pt/Fe heterogeneousparticulate catalyst versus the conversion of NPMIDA to glyphosate in acontinuous oxidation reactor system using a combination of Pt/Fe andPt/Fe/Te heterogeneous particulate catalysts.

The reactions were conducted in a continuous reactor system utilizing a2-liter Hastelloy C autoclave (Autoclave Engineers Inc., Pittsburgh,Pa.). The reactor was equipped with an agitator having a 1.25″ diametersix-blade turbine impeller, which was operated at 1600 RPM. The liquidlevel in the reactor was monitored using a Drexelbrook Universal III™Smart Level™, with a teflon-coated sensing element. An internal coolingcoil was utilized to control the temperature within the reactor duringthe course of the reaction.

In the first experiment, the reactor was loaded with a Pt/Feheterogenous particulate catalyst (2.18 g) and an aqueous slurry feedmaterial (1448 g). The catalyst comprised platinum (5% by weight) andiron (0.5% by weight). The aqueous slurry feed material comprised NPMIDA(3.5% by weight), glyphosate (1.5% by weight), formaldehyde (1200 ppm byweight), and formic acid (2500 ppm by weight). The slurry feed alsocontained NaCl (580 ppm by weight) to mimic NaCl impurity.

The reactor was pressurized to 100 psi with nitrogen and heated to 100°C. Once at temperature, a continuous flow of gaseous oxygen was fed tothe reactor without any liquid flow through the system. After 9 minutes,the continuous slurry feed was initiated at a rate of 70.4 g/min and aoxygen flow was continued as described in Table 49 below. A liquidproduct stream containing glyphosate product was continuously withdrawnfrom the reactor and analyzed by HPLC. Oxidation results are alsopresented in Table 49.

In the second experiment, the reactor was loaded with a Pt/Feheterogenous particulate catalyst (1.09 g), a Pt/Fe/Te heterogeneousparticulate catalyst (1.09 g) and an aqueous slurry feed material (1455g). The Pt/Fe catalyst comprised platinum (5% by weight) and iron (0.5%by weight) and the Pt/Fe/Te catalyst comprised platinum (5% by weight),iron (0.5% by weight) and tellurium (0.2% by weight). The aqueous slurryfeed material comprised NPMIDA (3.5% by weight), glyphosate (1.5% byweight), formaldehyde (1200 ppm by weight), and formic acid (2500 ppm byweight). The slurry feed also contained NaCl (580 ppm by weight) tomimic NaCl impurity.

The reactor was pressurized to 100 psi with nitrogen and heated to 100°C. Once at temperature, a continuous flow of gaseous oxygen was fed tothe reactor without any liquid flow through the system. After 19minutes, the continuous slurry feed was initiated at a rate of 70.4g/min and oxygen flow was continued as described in Table 50 below. Aliquid product stream containing glyphosate product was continuouslywithdrawn from the reactor and analyzed by HPLC. Oxidation results forthe second experiment are also presented in Table 50.

TABLE 49 Oxidation Results for Pt/Fe catalyst (Experiment 1) Elapsedfeed exit feed exit feed exit feed exit Time O₂ Flow NPMIDA NPMIDAGlyphosate Glyphosate CH₂O CH₂O HCOOH HCOOH (min) (sccm) (wt %) (wt %)(wt %) (wt %) (ppm) (ppm) (ppm) (ppm) 77.0 300.0 3.49 0.82 1.51 3.471098.9 2083.9 2118.7 4385.0 147.0 300.0 3.49 0.52 1.51 3.23 1098.91674.8 2118.7 4653.9 205.0 300.0 3.49 0.80 1.51 3.49 1098.9 2195.22118.7 4206.5 280.0 300.0 3.49 0.84 1.51 3.48 1098.9 2215.4 2118.74167.7 1212.0 300.0 3.49 1.07 1.51 3.40 1098.9 2344.2 2118.7 3991.11378.0 300.0 3.49 1.18 1.51 3.40 1098.9 2361.3 2118.7 3973.1 1447.0300.0 3.49 1.17 1.51 3.38 1098.9 2347.3 2118.7 4008.4 1618.0 300.0 3.491.26 1.51 3.35 1098.9 2323.1 2118.7 3985.8 1795.0 300.0 3.49 1.35 1.513.31 1098.9 2356.2 2118.7 3896.4 2683.0 300.0 3.49 1.39 1.51 3.27 1098.92316.2 2118.7 3861.0 2789.0 300.0 3.49 1.45 1.51 3.30 1098.9 2353.92118.7 3871.0 2885.0 300.0 3.49 1.53 1.51 3.25 1098.9 2310.5 2118.73796.1 3071.0 450.0 3.49 0.98 1.51 3.43 1098.9 2520.4 2118.7 3935.23180.0 700.0 3.49 0.88 1.51 3.55 1098.9 2653.0 2118.7 4086.1

TABLE 50 Oxidation Results for Pt/Fe and Pt/Fe/Te catalysts (Experiment2) feed exit feed exit feed exit feed exit Elapsed O₂ Flow NPMIDA NPMIDAGlyphosate Glyphosate CH₂O CH₂O HCOOH HCOOH Time (min) (sccm) (wt %) (wt%) (wt %) (wt %) (ppm) (ppm) (ppm) (ppm) 48.0 315.0 3.25 0.82 1.60 3.181222.1 940.17 2299.6 3757.3 156.0 330.0 3.25 0.75 1.60 3.15 1222.11087.2 2299.6 3975.6 210.0 365.0 3.25 0.57 1.60 3.27 1222.1 1200.92299.6 4114.4 281.0 410.0 3.25 0.46 1.60 3.39 1222.1 1306.9 2299.64182.9 339.0 410.0 3.25 0.67 1.60 3.35 1222.1 1306.5 2299.6 4191.2 626.0400.0 3.25 0.92 1.60 3.30 1222.1 1385.2 2299.6 4081.0 1295.0 425.0 3.251.10 1.60 3.15 1222.1 1289.0 2299.6 3910.2 1424.0 450.0 3.25 1.16 1.603.14 1222.1 1341.6 2299.6 3951.3 1548.0 450.0 3.25 1.13 1.60 3.07 1222.11292.0 2299.6 3948.4 1648.0 450.0 3.25 1.16 1.60 3.17 1222.1 1264.62299.6 3916.0 1762.0 450.0 3.25 1.26 1.60 3.11 1222.1 1234.7 2299.63964.7 1820.0 500.0 3.25 1.08 1.60 3.08 1222.1 1200.8 2299.6 4065.82749.0 500.0 3.25 1.78 1.60 2.76 1222.1 1079.1 2299.6 3927.5 2857.0500.0 3.25 1.92 1.60 2.75 1222.1 1065.1 2299.6 3926.3 2986.0 500.0 3.251.69 1.60 2.68 1222.1 1031.1 2299.6 3910.7 3118.0 500.0 3.25 1.81 1.602.64 1222.1 1009.7 2299.6 3892.2

The present invention is not limited to the above embodiments and can bevariously modified. The above description of the preferred embodiments,including the Examples, is intended only to acquaint others skilled inthe art with the invention, its principles, and its practicalapplication so that others skilled in the art may adapt and apply theinvention in its numerous forms, as may be best suited to therequirements of a particular use.

With reference to the use of the word(s) comprise or comprises orcomprising in this entire specification (including the claims below),Applicants note that unless the context requires otherwise, those wordsare used on the basis and clear understanding that they are to beinterpreted inclusively, rather than exclusively, and that Applicantsintend each of those words to be so interpreted in construing thisentire specification.

1. A process for removal of water from an aqueous starting solutioncomprising N-(phosphonomethyl)glycine product and crystallizationN-(phosphonomethyl)glycine product therefrom, the process comprising:introducing an aqueous evaporation feed mixture into an evaporation zoneof a crystallizer, said feed mixture comprising said aqueous startingsolution; evaporating water from said feed mixture in said evaporationzone in the presence of solid particulate N-(phosphonomethyl)glycineproduct, thereby producing a vapor phase comprising water vapor,precipitating N-(phosphonomethyl)glycine product from the aqueous liquidphase, and producing an evaporation product comprisingN-(phosphonomethyl)glycine product solids and a mother liquor that issubstantially saturated or supersaturated in N-(phosphonomethyl)glycineproduct; removing N-(phosphonomethyl)glycine product solids and motherliquor obtained in the evaporation product from the crystallizer; andmaintaining a ratio of particulate N-(phosphonomethyl)glycine productsolids to mother liquor in said evaporation zone which exceeds the ratioof N-(phosphonomethyl)glycine product solids incrementally produced bythe effects of evaporation to mother liquor incrementally producedthereby, wherein maintaining said ratio of particulateN-(phosphonomethyl)glycine product solids to mother liquor in saidevaporation zone comprises controlling the relative rates mother liquorand N-(phosphonomethyl)glycine product solids obtained in theevaporation product are removed from the crystallizer.
 2. The process asset forth in claim 1 wherein water is evaporated from said feed mixtureunder substantially adiabatic conditions in said evaporation zone of thecrystallizer.
 3. The process as set forth in claim 2 wherein saidevaporation product is divided to provide an N-(phosphonomethyl)glycineproduct solids fraction that is relatively depleted in mother liquor anda mother liquor fraction that is relatively depleted inN-(phosphonomethyl)glycine product solids.
 4. The process as set forthin claim 3 wherein maintaining said ratio of particulateN-(phosphonomethyl)glycine product solids to mother liquor in saidevaporation zone comprises returning solids obtained in said solidsfraction to said evaporation zone or retaining solids obtained in saidsolid fraction within said zone.
 5. The process as set forth in claim 4comprising: introducing an evaporation feed mixture comprising saidaqueous starting solution into a vapor/liquid separation zone of saidevaporation zone wherein the pressure is below the vapor pressure ofsaid mixture, thereby allowing water to flash from the evaporation feedmixture, producing a vapor phase comprising water vapor, andprecipitating N-(phosphonomethyl)glycine product from the liquid phaseto produce a first slurry stream comprising particulateN-(phosphonomethyl)glycine product in a saturated or supersaturatedmother liquor; separating said vapor phase from said first slurrystream; introducing said first slurry stream into a retention zone inwhich a supernatant liquid comprising a fraction of said mother liquoris separated from a second slurry stream comprising precipitatedN-(phosphonomethyl)glycine product and mother liquor, said retentionzone having an inlet for said first slurry, a decantation liquid exitfor said supernatant liquid spaced above said inlet, and an exit forsaid second slurry spaced above said inlet but below said decantationliquid exit; and maintaining the relative rates at which said firstslurry is introduced into said retention zone, said second slurry isdrawn off through said second slurry exit and said supernatant liquid isdrawn off through said decantation liquid exit such that the upward flowvelocity in a lower region of said retention zone below said secondslurry exit is sufficient to maintain precipitatedN-(phosphonomethyl)glycine product in suspension in the liquid phasewhile the upward flow velocity in an upper region of said retention zoneabove said second slurry exit is below the sedimentation velocity of atleast 80% by weight of the N-(phosphonomethyl)glycine product particlesin said lower region.
 6. The process as set forth in claim 5 wherein atleast a portion of said second slurry stream is recirculated to saidvapor/liquid separation zone.
 7. The process as set forth in claim 6wherein at least a portion of said second slurry stream and said aqueousstarting solution together comprise the evaporation feed mixtureintroduced into said vapor/liquid separation zone.
 8. The process as setforth in claim 7 wherein said aqueous starting solution and said secondslurry stream are mixed before introduction into said vapor/liquidseparation zone.
 9. The process as set forth in claim 7 wherein a thirdslurry stream is removed from said lower region of said zone.
 10. Theprocess as set forth in claim 9 wherein said third slurry stream isremoved from said lower region through a slurry exit separate from saidsecond slurry exit.
 11. The process as set forth in claim 9 wherein saidthird slurry stream is obtained by dividing said second slurry streaminto a recirculation stream and said third slurry stream.
 12. Theprocess as set forth in claim 9 wherein the relative rates of the flowof said aqueous starting solution to said vapor/liquid separation zone,recirculation of all or part of said second slurry stream to saidvapor/liquid separation zone, withdrawal of said supernatant liquid fromsaid decantation liquid exit, withdrawal of said third slurry streamfrom said lower region of said retention zone, and return to saidevaporation zone of any liquid or solids bearing streams from anysolids/liquid separations to which said third slurry may be subjected,are sufficient to establish a ratio of N-(phosphonomethyl)glycineproduct solids to mother liquor in said lower region of said zone thatis higher than the ratio of precipitated solidN-(phosphonomethyl)glycine product incrementally produced by the effectsof evaporation to mother liquor incrementally produced thereby.
 13. Theprocess as set forth in claim 12 wherein the relative flow rates of saidstreams are controlled so that the N-(phosphonomethyl)glycine productsolids concentration in said lower region of said zone is at least abouttwice the concentration of N-(phosphonomethyl)glycine product solids inthe mixture of such solids and mother liquor that is or would beproduced by flashing of said aqueous starting solution in saidvapor/liquid zone in the absence of said recirculated second slurrystream.
 14. The process as set forth in claim 13 wherein solids areremoved from said third slurry to produce a recycle liquid fractionwhich is recirculated to said vapor/liquid separation zone, whereby saidevaporation feed mixture further comprises said recycle liquid fraction.15. The process as set forth in claim 14 wherein both said aqueousstarting solution and said recycle liquid fraction are mixed with saidsecond slurry stream prior to introduction into said vapor/liquidseparation zone.
 16. The process as set forth in claim 15 wherein therelative flow rates of all of said streams, including said recycleliquid fraction, are controlled so that the solids content of the slurryin said lower region of said zone is at least about 12% by weight. 17.The process as set forth in claim 15 wherein the relative flow rates ofsaid streams are controlled so that solids removed from said thirdslurry have a median cube weighted particle size of at least about 200μm.
 18. The process as set forth in claim 15 wherein the relative flowrates of said streams are controlled so that the solids removed fromsaid third slurry have a BET surface area not greater than about 0.09m²/g.
 19. The process as set forth in claim 17 wherein the upward flowvelocity in said lower region is at least four times the sedimentationvelocity of at least 80% by weight of the solids contained therein, andthe upward flow velocity in the upper region of the zone is less thanone fourth the sedimentation velocity of at least 80% by weight of thesolids in said second slurry.
 20. The process as set forth in claim 13wherein said feed mixture comprises a slurry ofN-(phosphonomethyl)glycine product in an aqueous liquid phase that issupersaturated in N-(phosphonomethyl)glycine product, said feed mixtureflowing along a recirculation path between said second slurry exit andan inlet to said vapor/liquid separation zone without substantial axialback-mixing.
 21. The process as set forth in claim 20 wherein thesurface area of the solids contained in said feed mixture is sufficientto allow production of crystalline N-(phosphonomethyl)glycine producthaving a median cube weighted particle size of at least about 200 μm.22. The process as set forth in claim 21 wherein the maximum extent ofsupersaturation expressed as the difference between theN-(phosphonomethyl)glycine product concentration in the aqueous liquidphase at any location within said recirculation path and the saturationconcentration of N-(phosphonomethyl)glycine product in the aqueousliquid phase at such location is not greater than about 0.7% by weight,basis the aqueous liquid phase.
 23. The process as set forth in claim 21wherein the integrated average extent of supersaturation expressed asthe difference between the N-(phosphonomethyl)glycine productconcentration in the aqueous liquid phase and the saturationconcentration of N-(phosphonomethyl)glycine product in the aqueousliquid phase over said recirculation path is not greater than about 0.5%by weight, basis the aqueous liquid phase.
 24. The process as set forthin claim 21 wherein the integrated average extent of supersaturationexpressed as the difference between the N-(phosphonomethyl)glycineproduct concentration in the aqueous liquid phase and the saturationconcentration of N-(phosphonomethyl)glycine product in the aqueousliquid phase over said recirculation path is at least 0.2% lower thanthe extent of supersaturation required to provide the samecrystallization productivity per unit working volume of a referenceevaporator consisting of a fully back mixed evaporation zone in whichthe ratio of N-(phosphonomethyl)glycine product solids to mother liquoris equal to the ratio of N-(phosphonomethyl)glycine product solidsincrementally produced by the effects of the evaporation to motherliquor incrementally produced thereby.
 25. The process as set forth inclaim 13 wherein the rate of recirculation of said second slurry to saidvapor/liquid separation zone is at least about 20 times the rate ofwithdrawal of said supernatant liquid from said decantation liquid exit.26. The process as set forth in claim 5 wherein said vapor/liquidseparation zone is positioned above the interface between said lowerregion and said upper region of said retention zone and is segregatedfrom said upper region, said vapor/liquid separation zone being in fluidflow communication with the lower region of said retention zone via adraft tube for flow of said first slurry from said separation zone tosaid lower region.
 27. The process as set forth in claim 5 which isoperated substantially without transfer of heat to or from saidvapor/liquid separation zone, said retention zone, said feed mixture orsaid second slurry.
 28. The process as set forth in claim 5 whereinprecipitation of N-(phosphonomethyl)glycine product upon evaporationresults primarily from cooling of the aqueous liquid phase in theevaporation zone.
 29. The process as set forth in claim 5 comprising:removing an evaporation product slurry from said lower region of saidretention zone; subjecting said evaporation product slurry tosolids/liquid separation to provide an N-(phosphonomethyl)glycineproduct solids fraction that is relatively depleted in mother liquor anda mother liquor fraction that is relatively depleted inN-(phosphonomethyl)glycine product solids; and returning solids obtainedin said solids fraction to said evaporation zone or retaining solidsobtained in said solid fraction within said zone, thereby maintaining aratio of particulate N-(phosphonomethyl)glycine product solids to motherliquor in said evaporation zone which exceeds the ratio ofN-(phosphonomethyl)glycine product solids produced by the evaporation tomother liquor produced thereby.
 30. A method for removal of water froman aqueous starting solution comprising N-(phosphonomethyl)glycineproduct and crystallization of N-(phosphonomethyl)glycine producttherefrom, the process comprising: introducing an evaporation feedmixture comprising said aqueous starting solution into a vapor/liquidseparation zone wherein the pressure is below the vapor pressure of saidmixture, thereby allowing water to flash from the evaporation feedmixture, producing a vapor phase comprising water vapor and increasingthe concentration of N-(phosphonomethyl)glycine product in the remainingliquid phase to a concentration in excess of the solubility ofN-(phosphonomethyl)glycine product, whereby N-(phosphonomethyl)glycineproduct precipitates from the liquid phase to produce a first slurrystream comprising particulate N-(phosphonomethyl)glycine product in asaturated or supersaturated mother liquor; separating said vapor phasefrom said first slurry stream; introducing said first slurry stream intoa decantation zone in which a supernatant liquid comprising a fractionof said mother liquor is separated from a second slurry streamcomprising precipitated N-(phosphonomethyl)glycine product and motherliquor, said decantation zone having an inlet for said first slurry, adecantation liquid exit for said supernatant liquid spaced above saidinlet, and an exit for said second slurry vertically spaced above saidinlet but below said supernatant liquid exit; and maintaining therelative rates at which said first slurry is introduced into saiddecantation zone, said second slurry is drawn off through said secondslurry exit and said supernatant liquid is drawn off through saiddecantation liquid exit such that the upward flow velocity in a lowerregion of said decantation zone below said second slurry exit issufficient to maintain precipitated N-(phosphonomethyl)glycine productin suspension in the liquid phase while the upward flow velocity in anupper region of said decantation zone above said second slurry exit isbelow the sedimentation velocity of at least 80% by weight of theN-(phosphonomethyl)glycine product particles in said lower region.